Oxidative cracking of alkanes with fluidized vanadium catalyst

ABSTRACT

Fluidizable catalysts for the gas phase oxygen-free oxidative cracking of alkanes, such as hexane, to one or more olefins, such as ethylene, propylene, and/or butylene. The catalysts comprise 1-15% by weight per total catalyst weight of one or more vanadium oxides (VOx), such as V2O5. The catalysts are disposed on an alumina support that is modified with cerium to influence catalyst acidity and characteristics of lattice oxygen at the catalyst surface. Various methods of preparing and characterizing the catalyst as well as methods for the gas phase oxygen free oxidative cracking of alkanes, such as hexane, to one or more olefins, such as ethylene, propylene, and/or butylene with improved alkane conversion and olefins product selectivity are also disclosed.

BACKGROUND OF THE INVENTION Technical Field

The present disclosure relates to fluidizable vanadium basedVO_(x)/Ce-γ-Al₂O₃ catalysts and processes using the catalysts for thecracking of alkanes to olefins, such as hexane to propylene, in theabsence of gas phase oxygen.

Description of the Related Art

The “background” description provided herein is for the purpose ofgenerally presenting the context of the disclosure. Work of thepresently named inventors, to the extent it is described in thisbackground section, as well as aspects of the description which may nototherwise qualify as prior art at the time of filing, are neitherexpressly or impliedly admitted as prior art against the presentinvention.

Olefins such as ethylene and propylene are important feedstock in thechemical industry. They are used for the production of polyethylene,polypropylene and other industrial intermediates. Conventionally,olefins are obtained by steam cracking and catalytic conversion ofethanol. Steam cracking processes suffer from high production costs dueto the use of petroleum feedstock and the energy consumption in thefurnace [T. Ren, M. Patel, and K. Blok, “Olefins from conventional andheavy feedstocks: Energy use in steam cracking and alternativeprocesses,” Energy, vol. 31, pp. 425-451, 2006.]. Additionally, ethanecrackers as well as the retrofitting of naphtha crackers were developedafter the rise of shale oil technology over the past decades [N. Rahimiand R. Karimzadeh, “Catalytic cracking of hydrocarbons over modifiedZSM-5 zeolites to produce light olefins: A review,” Appl. Catal. A Gen.,vol. 398, no. 1-2, pp. 1-17, 2011; and B. Yilmaz and U. Müller,“Catalytic applications of zeolites in chemical industry,” Top. Catal.,vol. 52, pp. 888-895, 2009.—each incorporated herein by reference in itsentirety]. Furthermore, olefins are produced from methanol using zeoliteprocesses such as the methanol to olefins (MTO) process. The twocommonly used catalysts in this process are ZSM-5 (MFI-type) and SAPO-34(CHA-type). MFI is a medium pore sized alumina-silicate with tenmembered ring pores. MFI is highly selective toward propylene andbutylene; however, the yield of short olefins is less than in SAPO-34.SAPO-34 is a silico-alumino-phosphate with small eight-membered ringpores, which has high selectivity towards ethylene but suffers from fastdeactivation due to coke formation [D. Chen, K. Moljord, and a. Holmen,“A methanol to olefins review: Diffusion, coke formation anddeactivation on SAPO type catalysts,” Microporous Mesoporous Mater.,vol. 164, pp. 239-250, 2012; and S. Askari, R. Halladj, and M. Sohrabi,“Methanol conversion to light olefins over sonochemically preparedSAPO-34 nanocatalyst,” Microporous Mesoporous Mater., vol. 163, pp.334-342, 2012; and G. Liu, P. Tian, Q. Xia, and Z. Liu, “An effectiveroute to improve the catalytic performance of SAPO-34 in themethanol-to-olefin reaction,” J. Nat. Gas Chem., vol. 21, no. 4, pp.431-434, 2012.—each incorporated herein by reference in its entirety].Therefore, in general zeolites have deactivation problems and generallylow selectivity to olefins (i.e. less than 50%) [Mamedov, E. A.Corberfin, V. Cortds “Oxidative dehydrogenation of lower alkanes onvanadium oxide-based catalysts. The present state of the art andoutlooks” Appl. Catal. A Gen., vol. 127, pp. 1-40, 1995; and J. Li, Y.Wei, G. Liu, Y. Qi, P. Tian, B. Li, Y. He, and Z. Liu, “Comparativestudy of MTO conversion over SAPO-34, H-ZSM-5 and H-ZSM-22: Correlatingcatalytic performance and reaction mechanism to zeolite topology,”Catal. Today, vol. 171, no. 1, pp. 221-228, 2011; and J. Lefevere, S.Mullens, V. Meynen, and J. Van Noyen, “Structured catalysts formethanol-to-olefins conversion: a review,” Chem. Pap., vol. 68, no. 9,pp. 1143-1153, 2014.]. Catalytic cracking of hexane is another promisingprocess to produce short olefins, in these processes many types ofcatalyst have been used including metal oxides as well as zeolites.Zeolites have been tested for n-hexane cracking to propylene, such asthe zeolite MCM-22 with various Si/Al ratios [Y. Wang, T. Yokoi, S.Namba, J. N. Kondo, and T. Tatsumi, “Catalytic cracking of n-hexane forproducing propylene on MCM-22 zeolites,” Appl. Catal. A Gen.,2014.—incorporated herein by reference in its entirety]. Propyleneselectivity was found to be 40% at a Si/Al ratio of 62. ZSM-5 zeoliteswere also investigated using methanol coupling [F. Chang, Y. Wei, X.Liu, Y. Qi, D. Zhang, Y. He, and Z. Liu, “An improved catalytic crackingof n-hexane via methanol coupling reaction over HZSM-5 zeolitecatalysts,” Catal. Letters, vol. 106, no. February, pp. 171-176,2006.—incorporated herein by reference in its entirety]. Methanolcontributed positively to the reaction by decreasing the activationenergy which ultimately led to an increased olefins yield. Another studyinvestigated the cracking of hydrocarbons over a ZSM-5 catalyst withdifferent Si/Al ratios [B. Lücke, A. Martin, H. Günschel, and S. Nowak,“CMHC: coupled methanol hydrocarbon cracking,” Microporous MesoporousMater., vol. 29, pp. 145-157, 1999.—incorporated herein by reference inits entirety]. The catalyst was also modified using iron (Fe) todecrease catalyst deactivation; this increased the yield of olefins toas high as 305. In addition, metal oxides, such as MoO₂ were consideredas potential catalysts for n-hexane cracking obtaining a total olefinsyield of 85% [J. H. Song, P. Chen, S. H. Kim, G. a. Somorjai, R. J.Gartside, and F. M. Dautzenberg, “Catalytic cracking of n-hexane overMoO2,” J. Mol. Catal. A Chem., vol. 184, pp. 197-202, 2002.—incorporatedherein by reference in its entirety]. Catalytic oxidative cracking is apotential alternative to steam cracking due to factors including: (i) anexothermic oxidation reaction, (ii) an adiabatic process, and (iii) areduced coke formation. Although, for the non-catalytic pyrolysis ofhexane at 750° C. the oxygen feed was found to increase reaction rateand give a conversion of 85% with an olefin selectivity of 59% andethylene as the major product [Xiaoyin Chena, Yong Liua, Guoxing Niub,Zhuxian Yanga, Maiying Biana, Adi Hea, “High temperature thermalstabilization of alumina modified by lanthanum species” React. Kinet.Catal. Lett, vol. 81, no. 2, pp. 203-209, 2001.—incorporated herein byreference in its entirety]. The presence of oxygen allows the crackingto proceed in an autothermal way, where the exothermic reaction providesthe heat for the cracking reaction. Liu, et al. conducted acomprehensive study of homogeneous gas phase versus heterogeneouscatalytic oxidative cracking of hexane at a temperature of 700° C. [H.X. X. Liu, W. Li, H. Zhu, Q. Ge, Y. Chen, “Light Alkenes Preparation bythe Gas Phase Oxidative Cracking or Catalytic Oxidative Cracking of HighHydrocarbons”, Catal. Lett., vol. 94, no. 1-2, pp. 31-36,2004.—incorporated herein by reference in its entirety]. Amongst thecatalysts tested, 0.25 wt % Li/MgO showed the best performance (64 mol %conversion of hexane and 67 mol % selectivity to olefins). However, dueto the high reaction temperature, the gas phase reaction had the majorrole and the presence of catalyst had no major influence on conversionsof hexane and yields of olefins. Li/MgO catalysts with differentpromoters (i.e. MoO3, Bi2O3, V2O5) were tested for oxidative cracking ofn-hexane and the olefin selectivity obtained was up to 50% [C.Boyadjian, B. Van Der Veer, I. V. Babich, L. Lefferts, and K. Seshan,“Catalytic oxidative cracking as a route to olefins: Oxidativeconversion of hexane over MoO3-Li/MgO,” Catal. Today, vol. 157, pp.345-350, 2010.—incorporated herein by reference in its entirety].

Oxidative cracking of n-hexane over supported metal oxides is apotential synthetic route for the production of short olefins. Studieshave shown that the presence of metal oxide catalysts during crackingenhances yield of olefins. This observation can be explainedqualitatively with a mechanism that includes activation of the alkane onthe catalyst generating alkyl radicals that undergo a radical-chainmechanism in the gas phase. In this mechanism oxygen has two functions.First, the presence of small amounts of oxygen influences the radicalgas phase chemistry significantly because the type and concentration ofthe chain propagator radicals are greatly increased. At higher oxygenpartial pressures the radical chemistry is only slightly influenced bythe increasing oxygen concentration. The second function of oxygen is tofacilitate the removal of hydrogen from the surface OH-species that areformed during the activation of alkane on the catalyst. Therefore, ithas been shown that the oxidative conversion of propane over Li/MgOcatalysts follows a mixed heterogeneous-homogeneous radical chemistrywhere the catalyst acts as an initiator [L. Leveles, K. Seshan, J. A.Lercher, and L. Lefferts, “Oxidative conversion of propane overlithium-promoted magnesia catalyst—I. Kinetics and mechanism,” J.Catal., vol. 218, pp. 296-306, 2003.—incorporated herein by reference inits entirety]. Hence, the mentioned mechanism was used to study n-hexanecracking over Li/MgO modified catalysts, and it was proposed that hexaneis activated before the cracking reaction takes place. The study alsoshowed that the catalyst increases the production of olefins viadehydrogenation. Oxidative cracking of n-hexane to olefins is apromising process; however, all previous work was done using fixed-bedreactors and gas phase oxygen as a co-reactant for the purpose ofdehydrogenation. This setup induces CO_(x) formation by combustion dueto the presence of oxygen and therefore a lower yield of olefins isobtained.

In view of the forgoing, one aspect of the present invention is toprovide fluidizable oxidative cracking catalysts comprising vanadiumoxide catalytic species using a mixed Ce-γ-Al₂O₃ as support material.The physiochemical characterization of the catalysts offers anexamination of the VO_(x) monovanadate and polyvanadate surface specieson the support, the catalyst's stability, level of acidity andmetal-support interactions. A further aim of the present disclosure isto provide methods for producing these VO_(x)/Ce—Al₂O₃ catalysts. Anadditional aim of the present disclosure is to provide methods for theoxidative cracking of alkanes, such as hexane, to produce one or moreolefins, such as ethylene, propylene, and/or butylene employing thelattice oxygen of these VO_(x)/Ce—Al₂O₃ catalysts. These catalystspresent relatively high acidity enhancing the cracking reaction, andsimultaneously the catalyst lattice oxygen allows dehydrogenation ofalkanes to olefins to take place (FIG. 1). These methods may beperformed in a gas phase oxygen free environment under fluidized bedreaction conditions that enhance catalyst-feed contact at differenttemperatures and reaction times accomplishing high alkane conversion andhigh olefins product selectivity over CO_(x) combustion products.

BRIEF SUMMARY OF THE INVENTION

According to a first aspect, the present disclosure relates to acatalyst comprising i) a support material comprising alumina modified bycerium and ii) a catalytic material comprising one or more vanadiumoxides disposed on the support material, wherein the catalyst comprises1-15% of the one or more vanadium oxides by weight relative to the totalweight of the catalyst.

In one embodiment, the catalyst comprises 0.05-1.0% of cerium by weightrelative to the total weight of the catalyst.

In one embodiment, the one or more vanadium oxides form an amorphousphase on the surface of the support material.

In one embodiment, the one or more vanadium oxides form a crystallinephase on the surface of the support material.

In one embodiment, the one or more vanadium oxides are at least oneselected from the group consisting of V₂O₅, VO₂, and V₂O₃.

In one embodiment, the catalyst comprises at least 50% of V₂O₅ by weightrelative to the total weight of the one or more vanadium oxides.

In one embodiment, the catalyst has an average particle size in therange of 20-160 μm.

In one embodiment, the catalyst has an apparent particle density in therange of 1-10 g/cm³.

In one embodiment, the catalyst has a BET surface area in the range of2-50 m²/g.

In one embodiment, the catalyst is fluidizable and has Class B powderproperties in accordance with Geldart particle classification.

According to a second aspect, the present disclosure relates to a methodfor producing the catalyst of the present disclosure in any of itsembodiments comprising i) mixing an aluminum salt or hydrate with acerium salt or hydrate in a solvent to form an alumina precursorsolution, ii) adding a base to and hydrolyzing the alumina precursorsolution to form the support material comprising alumina modified bycerium, iii) mixing the support material with a solution comprising avanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors, iv) reducing the loaded catalyst precursors with H₂gas to form reduced catalyst precursors, and v) oxidizing the reducedcatalyst precursors with oxygen to form the catalyst.

According to a third aspect, the present disclosure relates to a methodfor the oxidative cracking of an alkane to produce one or more olefinscomprising flowing the alkane through a reactor comprising a catalystchamber loaded with the catalyst of claim 1 at a temperature in therange of 450-700° C. to form the one or more olefins and a reducedcatalyst.

In one embodiment, the reactor is a fluidized bed reactor and theoxidative cracking is performed in a gas phase oxygen free environment.

In one embodiment, the alkane is hexane the one or more olefins compriseat least one of ethylene, propylene, butylene, and mixtures thereof.

In one embodiment, the method further comprises further comprises i)oxidizing at least a portion of the reduced catalyst by flowing airthrough the catalyst chamber to regenerate the catalyst and ii)repeating the flowing and the oxidizing at least once with a less than15% decrease in percent conversion of the alkane, a less than 15%decrease in selectivity for the one or more olefins relative to a totalpercentage of products formed, or both.

In one embodiment, the catalyst is present at an amount in the range of0.50-2.5 g of catalyst per mL of alkane.

In one embodiment, the alkane is hexane and the method has a hexaneconversion of 5-50 mol % at a reaction time of 1-40 seconds and atemperature of 500-650° C.

In one embodiment, the alkane is hexane and the method has a shortolefins selectivity defined as moles of ethylene, propylene, andbutylene produced per moles of hexane converted of 20-70% at a reactiontime of 1-40 seconds and a temperature of 500-650° C.

In one embodiment, the alkane is hexane and the method has a CO_(x)selectivity defined as moles of carbon monoxide and carbon dioxideproduced per moles of hexane converted of no more than 50% at a reactiontime of 1-40 seconds and a temperature of 500-650° C.

In one embodiment, the alkane is hexane and the method has a shortolefins selectivity defined as moles of ethylene, propylene, andbutylene produced per moles of hexane converted of at least 60% at areaction time of 1-40 seconds and a temperature of 500-650° C.

The foregoing paragraphs have been provided by way of generalintroduction, and are not intended to limit the scope of the followingclaims. The described embodiments, together with further advantages,will be best understood by reference to the following detaileddescription taken in conjunction with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

A more complete appreciation of the disclosure and many of the attendantadvantages thereof will be readily obtained as the same becomes betterunderstood by reference to the following detailed description whenconsidered in connection with the accompanying drawings, wherein:

FIG. 1 is an exemplary schematic representation of the oxidativecracking reaction of hexane by the prepared VO_(x)/Ce-γ-Al₂O₃ catalystsample in a gas phase oxygen free environment.

FIG. 2 is the X-ray diffraction (XRD) patters of the preparedVO_(x)/Ce-γ-Al₂O₃ catalyst sample and the Ce-γ-Al₂O₃ support.

FIG. 3 is the Fourier transform infrared (FTIR) absorption spectra ofthe prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample.

FIG. 4 is the laser Raman spectroscopy spectra of the preparedVO_(x)/Ce-γ-Al₂O₃ catalyst sample.

FIG. 5 is the temperature programmed reduction-oxidation (TPR/TPO)profiles of the prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample over multiplecycles.

FIG. 6 is the ammonia temperature programmed desorption (NH₃-TPD)profile of the prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample withtemperature rising gradually with a constant value β=10° C./min anddeconvolution peaks.

FIG. 7 is a comparison of the experimental data and fitted model ofammonia desorption during NH₃-TPD kinetics analysis for the preparedVO_(x)/Ce-γ-Al₂O₃ catalyst sample.

FIG. 8 is a graph of hexane (C₆H₁₄) conversion and CO_(x), CH₄,parrafins (C₂ to C₅) and olefins (C₂H₄, C₃H₆, and C₄H₈) productselectivity for the prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample in theoxidative cracking of hexane reaction runs at a temperature range of525-600° C., a reaction time of 15 seconds, and a hexane feed of 0.4 mLat STP illustrating the effect of temperature on the hexane conversionand product selectivity.

FIG. 9 is a graph of hexane (C₆H₁₄) conversion and CO_(x), CH₄,parrafins (C₂ to C₅) and olefins (C₂H₄, C₃H₆, and C₄H₈) productselectivity for the prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample in theoxidative cracking of hexane reaction runs at a reaction time range of5-25 seconds, a reaction temperature of 550° C., and a hexane feed of0.4 mL at STP illustrating the effect of reaction time on the hexaneconversion and product selectivity.

FIG. 10 is graph of hexane (C₆H₁₄) conversion and CO_(x), CH₄, parrafins(C₂ to C₅) and olefins (C₂H₄, C₃H₆, and C₄H₈) product selectivity forthe prepared VO_(x)/Ce-γ-Al₂O₃ catalyst sample without regeneration inthe oxidative cracking of hexane reaction runs at reaction temperaturesof 550° C. and 575° C., a reaction time of 15 seconds, and a hexane feedof 0.4 mL at STP illustrating the effect of lacking surface oxygen onthe hexane conversion and products selectivity.

DETAILED DESCRIPTION OF THE EMBODIMENTS

Referring now to the drawings. Embodiments of the present disclosurewill now be described more fully hereinafter with reference to theaccompanying drawings, in which some, but not all of the embodiments ofthe disclosure are shown.

According to a first aspect, the present disclosure relates to acatalyst, comprising i) a support material comprising alumina modifiedby cerium and ii) a catalytic material comprising one or more vanadiumoxides disposed on the support material, wherein the catalyst comprises1-15% of the one or more vanadium oxides by weight relative to the totalweight of the catalyst.

Vanadium oxide (e.g., V₂O₅—vanadia) is considered to be one of the mostimportant and useful metals to be used as a catalyst due to its physicaland chemical properties, and catalysis is the most dominantnon-metallurgical use of vanadia. The catalytic activity of vanadia isattributed to its reducible nature and its ability to easily change itsoxidation state from V⁺³ to V⁺⁵. It is generally accepted that V⁺⁵ isthe highly active initial state of the catalyst in a cycle of oxidativedehydrogenation. Vanadium oxide catalysts have been used in manyindustrial and lab scale catalytic reactions and processes. In manycases, vanadia catalysts are doped with promoters to improve theiractivity or selectivity, while various supports are used to improvemechanical strength, thermal stability, longevity, and/or catalyticperformance.

As used herein, a catalyst support material refers to material, usuallya solid with a high surface area, to which a catalyst is affixed. Thereactivity of heterogeneous catalyst and nanomaterial based catalystsoccurs at the surface atoms. Thus, great effort is made herein tomaximize the surface of a catalyst by distributing it over the support.The support may be inert or participate in the catalytic reactions. Thesupport materials used in catalyst preparation play a role indetermining the physical characteristics and performance of thecatalysts. Typical supports include various kinds of carbon, alumina andsilica. In a preferred embodiment, the dehydrogenation catalyst of thepresent disclosure comprises an alumina support material, preferably acerium modified alumina support material.

As used herein, alumina refers to aluminum oxide, a chemical compound ofaluminum and oxygen with the chemical formula Al₂O₃. Aluminum oxide iscommonly called alumina and may also be referred to as aloxide, aloxite,or alundum. It is the most commonly occurring of several aluminum oxidesand specifically identified as aluminum (III) oxide. It commonly occursin its crystalline polymorphic phase α-Al₂O₃ which composes the mineralcorundum, the most thermodynamically stable form of aluminum oxide.Al₂O₃ is significant in its use to produce aluminum metals and noted forits high melting point. In one embodiment, the catalytic material isloaded on an inert alumina support. Exemplary inert alumina based inertmaterials include, but are not limited to aluminum oxide, alumina,alumina monohydrate, alumina trihydrate, alumina silica, bauxite,calcined aluminum hydroxides such as gibbsite, bayerite and boehmite aswell as calcined hydrotalcite and the like.

In one embodiment, the alumina support material may be comprised of aplurality of different crystallographic phases. In the most common andthermodynamically stable form, corundum, the oxygen ions nearly form ahexagonal close-packed structure with aluminum ions filling two-thirdsof the octahedral interstices. Each Al³⁺ center is octahedral. In termof its crystallography, corundum adopts a trigonal Bravais lattice andits primitive cell contains two formula units of aluminum oxide.Aluminum oxide also exists in other phases, including the transitioncubic γ and η phases, the monoclinic θ phase, the hexagonal χ phase, theorthorhombic κ phase and the transition δ phase that can be tetragonalor orthorhombic. Each has unique crystal structure and properties. Inthe present disclosure, aluminum oxide or alumina may refer to Al₂O₃having an α polymorph, a γ polymorph, a η polymorph, a θ polymorph, a χpolymorph, a κ polymorph and a δ polymorph or mixtures thereof,preferably a γ polymorph. In a preferred embodiment, the alumina of thepresent disclosure consists substantially of γ-Al₂O₃, preferably greaterthan 75% by weight relative to the total weight of alumina, preferablygreater than 80%, preferably greater than 85%, preferably greater than90%, preferably greater than 95%, preferably greater than 98%,preferably greater than 99% by weight relative to the total weight ofthe alumina. In at least one embodiment, the alumina support materialconsists essentially of γ-alumina (γ-Al₂O₃).

Alumina, especially γ-Al₂O₃ is used for its very high surface area onwhich active metal atoms/crystallites can spread out as reactive sites,but also for its enhancement of productivity and/or selectivity throughmetal-support interaction and spillover/reverse-spillover phenomena. Inreactions, γ-Al₂O₃ must retain as much high surface area duringreaction. Additives and/or modifiers and additional supports markedlyincrease the thermal stability of the support, effect acidity and activesite metal support-interactions and prevent the loss of surface areaunder thermal reaction conditions.

In a preferred embodiment, the support material comprising alumina ismodified by cerium metal. As used herein, cerium refers to the chemicalelement with symbol Ce and atomic number 58. Cerium is the most abundantof the rare earth elements or rare earth metals and is a soft silveryductile metal belonging to the lanthanide group. Cerium has a variableelectronic structure and four allotropic forms of cerium are known toexist at standard pressure. The high temperature form, δ-cerium, has abcc (body-centered cubic) crystal structure and exists above 726° C. Thestable form below 726° C. to approximately room temperature is γ-cerium,with an fcc (face-centered cubic) crystal structure. The dhcp (doublehexagonal close-packed) form of β-cerium is the equilibrium structureapproximately from room temperature to −150° C. The fcc α-cerium existsbelow about −150° C. Other solid phases occurring at high pressures alsoexist. Both γ and β forms are quite stable at room temperature, althoughthe equilibrium transformation temperature is estimated at ˜75° C. Inthe present disclosure, cerium may refer to Ce having an α polymorph, aγ polymorph, a β polymorph, and a δ polymorph or mixtures thereof,preferably a γ or β polymorph.

In a preferred embodiment, the catalyst of the present disclosurecomprises 80-99% of alumina by weight relative to the total weight ofthe catalyst, preferably 82-98%, preferably 84-96%, preferably 86-95%,preferably 88-95% of alumina by weight relative to the total weight ofthe catalyst. In a preferred embodiment, the catalyst of the presentdisclosure comprises 0.05-1.0% of cerium by weight relative to the totalweight of the catalyst, preferably 0.1-0.8%, preferably 0.15-0.6%,preferably 0.16-0.4%, preferably 0.18-0.3%, or about 0.2% of cerium byweight relative to the total weight of the catalyst. In a preferredembodiment, the catalyst of the present disclosure has a weight ratio ofcerium to alumina in the range of 1:10000 to 1:100, preferably 1:8000 to1:200, preferably 1:6000 to 1:400, preferably 1:4000 to 1:600,preferably 1:2000 to 1:800, preferably 1:1500 to 1:1000.

It is equally envisaged that the catalyst of the present disclosure maybe adapted to incorporate additional rare earth metals. In someembodiments, these additional rare earth metals may be used in additionto, or in lieu of cerium. Exemplary additional rare earth metalsinclude, but are not limited to, light rare earth elements (LREE such asSc, La, Ce, Pr, Nd, Pm, Sm, Eu, and Gd; also known as the cerium group),and heavy rare earth elements (Y, Tb, Dy, Ho, Er, Tm, Yb, and Lu; alsoknown as the yttrium group). It is equally envisaged that the catalystof the present disclosure may be adapted to incorporate additionalsupport materials and additional additives such as phase transformationstabilizers. In some embodiments, these additional support materials andadditional additives may be used in addition to, or in lieu of aluminaand/or cerium. Exemplary additional support materials include, but arenot limited to oxides such as, SiO₂, TiO₂, ZrO₂, CeO, NbOs, MgO, CaO andzeolites. Exemplary additional thermal stabilizer additives include, butare not limited to, the elements La, Ce, Ba, Sr, Sm, Si, Pr and P. Whenlanthanum is used as an additive, the formation of lanthanum aluminatecan decrease the surface energies of γ-Al₂O₃ lowering the driving forcefor sintering and stabilizing bulk phase transformation. In certainembodiments, the catalyst of the present disclosure comprises less than5% of additional additives, such as elemental lanthanum, by weightrelative to the total weight of the catalyst, preferably 0.1-3.0% ofadditional additives by weight relative to the total weight of thecatalyst, preferably 0.5-2.0%, preferably 0.75-1.5%, preferably0.8-1.1%, or about 1.0% of additional additives by weight relative tothe total weight of the catalyst.

In a preferred embodiment, the catalyst of the present disclosurecomprises a catalytic material disposed on the support material, whereinthe catalytic material comprises one or more vanadium oxides. As usedherein, “disposed on” or “impregnated” describes being completely orpartially filled throughout, saturated, permeated and/or infused. Thecatalytic material may be affixed on one or more surfaces of the supportmaterial the catalytic material may be affixed on an outer surface ofthe support material or within pore spaces of the support material. Thecatalytic material may be affixed to the support material in anyreasonable manner, such as physisorption or chemisorption and mixturesthereof. In one embodiment, greater than 10% of the surface area (i.e.surface and pore spaces) of the support material is covered by thecatalytic material, preferably greater than 15%, preferably greater than20%, preferably greater than 25%, preferably greater than 30%,preferably greater than 35%, preferably greater than 40%, preferablygreater than 45%, preferably greater than 50%, preferably greater than55%, preferably greater than 60%, preferably greater than 65%,preferably greater than 70%, preferably greater than 75%, preferablygreater than 80%, preferably greater than 85%, preferably greater than90%, preferably greater than 95%, preferably greater than 96%,preferably greater than 97%, preferably greater than 98%, preferablygreater than 99%. In preferred embodiments, the vanadium or vanadiumoxide comprising catalytic material is homogeneously distributed ordispersed throughout the support material and on the surface of thesupport material. In preferred embodiments, this quality of thedispersion can be verified via scanning electron microscopy (SEM) and/orenergy dispersive X-ray analysis (EDX) providing elemental mapping,preferably vanadium elemental mapping. In other embodiments thecatalytic material may form localized clusters amongst the supportmaterial, form oxide species with the support catalyst or form layers ofthe catalytic material and vanadium species amongst the supportmaterial, or be heterogeneously disposed on the support material and itssurfaces and mixtures thereof.

In a preferred embodiment, the catalytic material comprises one or morevanadium oxides. In terms of the present disclosure, vanadium oxide mayrefer to vanadium (II) oxide (vanadium monoxide, VO), vanadium (III)oxide (vanadium sesquioxide or trioxide, V₂O₃), vanadium (IV) oxide(vanadium dioxide, VO₂), vanadium (V) oxide (vanadium pentoxide, V₂O₅).Vanadium oxide may also refer to a vanadate, a compound containing onoxoanion of vanadium generally in its highest oxidation state of ⁺5. Thesimplest vanadate ion is the tetrahedral orthovanadate VO₄ ³⁻ anion.Exemplary vanadate ions include, but are not limited to, VO₄ ³⁻, V₂O₇⁴⁻, V₃O₉ ³⁻, V₄O₁₄ ³⁻, V₅O₁₄ ³⁻ and the like. In addition to theseprincipal oxides of vanadium, various other distinct phases exist.Phases with the general formula V_(n)O_(2n+1), wherein n is a wholenumber greater than zero exist between V₂O₅ (vanadium (V) species) andvanadium (IV) species. Examples of these phases include V₃O₇, V₄O₉ andV₆O₁₃. Phases with the general formula V_(n)O_(2n+1), wherein n is awhole number greater than zero exist between vanadium (IV) species andV₂O₃ (vanadium (III) species). Termed Magneli phases, they are examplesof crystallographic shear compounds based on rutile structure. Examplesof Magneli phases include V₄O₇, V₅O₉, V₆O₁₁, V₇O₁₃ and V₈O₁₅. Manyvanadium oxygen phases are non-stoichiometric. In a preferredembodiment, the catalyst of the present disclosure comprises 1-15% ofthe one or more vanadium oxides by weight relative to the total weightof the catalyst, preferably 2-14%, preferably 3-12%, preferably 3.5-10%,preferably 4-8%, preferably 4.5-6%, or about 5% of the one or morevanadium oxides by weight relative to the total weight of the catalyst.

In a preferred embodiment, the one or more vanadium oxides are of theformula V_(x)O_(y) wherein x=1-4, preferably 1-3, more preferably 1-2and y=2-10, preferably 2-5. In a preferred embodiment, the one or morevanadium oxides are at least one selected from the group consisting ofV₂O₅, VO₂ and V₂O₃. V₂O₅ or vanadium (V) oxide or vanadium pentoxide isan inorganic compound that due to its high oxidation state is both anamphoteric oxide and an oxidizing agent. V₂O₅ is characterized by itsvaluable redox properties as V₂O₅ is easily reduced to the stablevanadium (IV) species. In certain embodiments, the catalyst comprises atleast 50% of V₂O₅ by weight relative to the total weight of the one ormore vanadium oxides, preferably greater than 60%, preferably greaterthan 70%, preferably greater than 80%, preferably greater than 85%,preferably greater than 90%, preferably greater than 95%, preferablygreater than 96%, preferably greater than 97%, preferably greater than98%, preferably greater than 99% of V₂O₅ by weight relative to the totalweight of the one or more vanadium oxides, such as, for example 50-90%by weight V₂O₅, preferably 75-80% V₂O₅, more preferably 85-90% V₂O₅,even more preferably at least 90-95% V₂O₅, most preferably 95-99.9% V₂O₅relative to the total weight of the one or more vanadium oxides. Incertain embodiments, the catalyst of the present disclosure consistsessentially of V₂O₅ and is substantially free of V₂O₃ and VO₂. In someembodiments, the catalyst of the present disclosure is substantiallyfree of V₂O₃ and comprises a mixture of at least 50% V₂O₅ by weightrelative to the total weight of the one or more vanadium oxides, withthe balance substantially comprising VO₂.

The different vanadia phases that can be present in supported vanadiaoxide catalysts as well as the distribution among the various vanadiumoxide structures can depend on the synthesis method, the vanadiumprecursor, solvent, calcination temperature, vanadium oxide loading,oxide support, etc. At loadings below “monolayer coverage” isolated andoligomerized surface VO₄ species may be present on the oxide support.The surface VO₄ species may possess up to three different oxygen atomsincluding, but not limited to, oxygen atoms forming a vanadyl group(V═O), oxygen atoms bridging two vanadia atoms (V—O—V), and oxygen atomsbridging a vanadia atom and oxide support cation (V—O-support).Depending on the vanadia surface density as well as the supportmaterial, a vanadia “monolayer coverage” may be reached. A “monolayer”refers to a single, closely packed layer of atoms or molecules, here theone or more vanadium oxides. As used herein, “monolayer coverage” refersto the completion of a 2D surface of vanadium oxide overlayer on thealumina support, and the surface becomes saturated before 3D vanadiumoxide and/or V₂O₅ crystallites start to form and grow subsequently. Incertain preferred embodiments, the vanadium loading is below themonolayer coverage and the VO_(x) species in the catalytic material arehighly dispersed forming an amorphous phase on the γ-Al₂O₃ and Cesupport surface. Alternatively, the monolayer coverage may be thought ofas the minimum amount of single vanadium and/or vanadium oxide atoms ormolecules to cover exactly 100% of the surface area (surface and porespaces) of the support material uniformly. In a preferred embodiment,the monolayer coverage of the dehydrogenation catalyst of the presentdisclosure corresponds to 5-20 vanadium atoms per nm² of support,preferably 6-15 atoms/nm², preferably 7-10 atoms/nm², preferably 8-9vanadium atoms per nm² of support. In certain embodiments, V₂O₅crystallites may be present at vanadium oxide loadings below monolayercoverage when a precursor vanadium salt is not well dispersed over thesupport during synthesis or when a weak interaction exists between thevanadium oxide and the support. In certain preferred embodiments, theone or more vanadium oxides may form a crystalline phase on the surfaceof the cerium modified alumina support material, preferably a V₂O₅crystalline phase. At high enough loadings, greater than monolayercoverage, vanadium oxide nanocrystals or nanoparticles having an averageparticle size of 1-100 nm, preferably 4-80 nm, preferably 10-60 nm,preferably 20-40 nm may be present on the surface of the catalystsupport. In certain embodiments, the different surface vanadia speciesmay be identified by techniques including, but not limited to, Ramanspectroscopy, Fourier transform infrared spectroscopy (FT-IR), UV-visspectroscopy, X-ray powder diffraction (XRD) and the like. In apreferred embodiment, the one or more vanadium oxides form an amorphousphase on the surface of the support material. In a preferred embodiment,the catalytic material comprising one or more vanadium oxides of thepresent disclosure forms a crystalline phase on the support surface. Inother embodiments, the catalytic material may display an amorphousphase, a crystalline phase, or both in the form of a mixed amorphous andcrystalline phase. In certain embodiments, the amorphous ornon-crystalline form of the one or more vanadium oxide species may befavorable for olefin selectivity, in other embodiments the crystallineform of the one or more vanadium oxides may be favorable for CO₂formation.

In certain embodiments, the catalytic material comprises one or morevanadium oxides and may optionally further comprise a promoter. As usedherein, a promoter refers to an additive to improve catalystperformance. Metal promoters such as for example niobium may function toisolate active species (i.e. VO_(x), more preferably V₂O₅) and to formsecondary metallic oxides (i.e. Nb₂O₅) on support surface. Furthermore,the addition of promoters to the catalytic material blocks acid siteswhich decreases the total acidity of the catalyst. In certainembodiments, the decrease in acidity and increase in basicity mayfacilitate desorption of substrates from the catalyst surface,preventing further oxidation, such as, for example the undesirablecombustion to carbon oxides (CO_(x)) in the oxidative dehydrogenation ofalkanes or oxidative cracking of alkanes such as hexane. In a preferredembodiment, the catalyst of the present disclosure may further comprise1.0-5.0% of promoter by weight relative to the total weight of thecatalyst, preferably 1.5-4.0%, preferably 2.0-3.75%, preferably3.0-3.5%, or about 3.25% of promoter by weight relative to the totalweight of the catalyst. Exemplary promoters include, but are not limitedto, metallic promoters (Nb, Cr, Mo, Ta, W), alkali promoters (Li, K, Rb)and halide promoters (Cl) and mixtures thereof. In preferredembodiments, the vanadium or vanadium oxide and promoter or promotersare homogeneously distributed throughout the catalyst support. In otherembodiments the promoter may form localized clusters amongst thevanadium, form promoter oxide species with the support catalyst, formlayers of promoter and vanadium species, or be disposed on the vanadiumoxide species and mixtures thereof.

In a preferred embodiment, the present disclosure provides fluidizablecatalysts for oxidative cracking and/or oxidative dehydrogenation (ODH)of alkanes preferably in reactors having a fluidized bed design. As usedherein “fluidizable” refers to the ability to undergo fluidization whichrefers to a process similar to liquefaction whereby a granular materialis converted from a static solid-like to a dynamic fluid-like state. Theprocess occurs when a fluid (liquid or gas) is passed up through thegranular material. A fluidized bed is formed when a quantity of a solidparticulate substance is placed under appropriate conditions to cause asolid/fluid mixture to behave as a fluid. This is usually achieved bythe introduction of pressurized fluid through the particulate medium.This results in the medium then having many properties andcharacteristics of normal fluids, such as the ability to free flow undergravity, or to be pumped using fluid type technologies. Fluidized bedtypes can be broadly classified by their flow behavior including, butnot limited to, stationary or bubbling fluidized beds, circulatingfluidized beds (CFB), vibratory fluidized beds, transport or flashreactor (FR), and annular fluidized beds (AFB).

In certain fluidized bed reactors, the catalyst pellets lie on a grateat the bottom of the reactor. Reactants are continuously pumped into thereactor through a distributor causing the bed to become fluidized.During the fluidization, the catalyst pellets are converted from astatic solid like state to a dynamic fluid like state. The bed'sbehavior after initial fluidization depends on the state of thereactant. If it is a liquid the bed expands uniformly with an increasedupward flow of the reactant, resulting in a homogeneous fluidization. Ifthe reactant is a gas, the bed will be non-uniform because the gas formsbubbles in the bed, resulting in aggregative fluidization. In terms ofthe present disclosure, the fluidization may be homogeneous oraggregative. In certain embodiments, the reactant or feed is preferablyan alkane including, but not limited to, ethane, propane, butane(including n-butane and isobutene), and hexane (including n-hexane,isohexanes, and neohexane) all of which may be present as gases andhence, an aggregative fluidization may be probable.

Properties and parameters for determining the fluidizability,reducibility, and oxygen carrying capacity of a catalyst can be bothmeasured and calculated. The average particle size and the particle sizedistribution can be measured, for example, using a Mastersizer 2000 fromMalvern Instruments. For spherical or substantially spherical catalystparticles, average particle size refers to the longest linear diameterof the catalyst particles. In a preferred embodiment, the catalyst ofthe present disclosure in any of its embodiments has an average particlesize in the range of 20-160 μm, preferably 30-150 μm, preferably 40-120μm, preferably 50-100 μm, more preferably 60-80 μm. In one embodiment,the particle size distribution of the catalyst of the present disclosureis 10-200 μm and greater than 75% of the particles have a particle sizeof 40-120 μm, preferably greater than 80%, preferably greater than 85%,more preferably greater than 90% have a particle size of 40-120 μm. Inanother embodiment, the catalyst of the present disclosure has aparticle size distribution ranging from 33% of the average particle sizeto 133% of the average particle size, preferably 50-130%, preferably60-125%, preferably 80-100%, preferably 90-110%, preferably 95-105% ofthe average particle size. In one embodiment, the catalyst particles ofthe present disclosure are monodisperse, having a coefficient ofvariation or relative standard deviation, expressed as a percentage anddefined as the ratio of the particle size standard deviation (σ) to theparticle mean size (μ) multiplied by 100 of less than 25%, preferablyless than 20%, preferably less than 15%, preferably less than 12%,preferably less than 10%, preferably less than 8%, preferably less than6%, preferably less than 5%.

As used herein, the apparent particle density refers to the mass of thecatalyst divided by the volume that it occupies. The apparent particledensity can be assessed using a CREC-established method. In the method,a known amount of catalyst is introduced to a flask. The flask is filledwith isopropanol and the apparent particle density (AD) is calculatedusing the following equation formula (I).

$\begin{matrix}{{A\; D} = \frac{W_{cat}}{V_{T} - V_{isopropanol}}} & (I)\end{matrix}$

Where AD is the apparent particle density (g/cm³), W_(cat) is thecatalyst weight, V_(T) is the flask volume and V_(isopropanol) is thevolume of isopropanol calculated as the ratio of the weight ofisopropanol needed to fill the flask and the density of isopropanol. Ina preferred embodiment, the catalyst of the present disclosure in any ofits embodiments has an apparent particle density of 1.0-10.0 g/cm³,1.1-5.0 g/cm³, preferably 1.25-4.0 g/cm³, preferably 1.5-3.5 g/cm³, morepreferably 1.8-3.2 g/cm³.

In some embodiments, with the calculated average particle size andparticle apparent density values, the fluidization regime of thecatalyst particles of the present disclosure can be determined usingGeldart's powder classification chart. Geldart groups powders into four“Geldart Groups” or “Geldart Classes”. The groups are defined bysolid-fluid density difference and particle size. Design methods forfluidized beds can be tailored based upon a particle's Geldart Group.For Geldart Group A the particle size is between 20 and 100 μm and theparticle density is typically less than 1.4 g/cm³. Prior to theinitiation of a bubbling bed phase, beds from these particles willexpand by a factor of 2 to 3 at incipient fluidization, due to tadecreased bulk density. Most powder-catalyzed beds utilize this group.For Geldart Group B the particle size lies between 40 and 500 μm and theparticle density is between 1.4-4 g/cm³. Bubbling typically formsdirectly at incipient fluidization. For Geldart Group C the groupcontains extremely fine and consequently the most cohesive particles.With a particle size of 20 to 30 μm, these particles fluidize under verydifficult to achieve conditions, and may require the application of anexternal force, such as mechanical agitation. For Geldart Group D theparticles in this regime are above 600 μm and typically have highparticle densities. Fluidization of this group requires very high fluidenergies and is typically associated with high levels of abrasion.Additionally, these particles are usually processed in shallow beds orin the spouting mode. The catalyst of the present disclosure ispreferably fluidizable and may be classified as a Geldart Group Apowder, a Geldart Group B powder, a Geldart Group C powder or a GeldartGroup D powder, preferably as a Geldart Group B powder. In at least onepreferred embodiment, the catalyst particles display a Geldart Group Bpowder property, which is highly fluidizable under oxidative crackingand ODH conditions. Large particles, such as those under Geldart GroupD, may limit the gas phase reactant access to the inner layers of thecatalyst. As a result, using smaller particles can minimize thediffusional resistance and reduction/oxidation rates can be maximized.On the other hand, very small particles, such as those under Geldart'sGroup C, can cause fluidization problems, channeling and loss of fines.

The Brunauer-Emmet-Teller (BET) theory aims to explain the physicaladsorption of gas molecules on a solid surface and serves as the basisfor an important analysis technique for the measurement of the specificsurface area of a material. Specific surface area is a property ofsolids which is the total surface area of a material per unit of mass,solid or bulk volume, or cross sectional area. In a preferredembodiment, the catalyst of the present disclosure in any of itsembodiments has a BET surface area in the range of 25-400 m²/g,preferably 50-350 m²/g, preferably 75-300 m²/g, preferably 100-250 m²/g,preferably 125-225 m²/g, preferably 150-200 m²/g.

The catalytic activity of many oxides in various processes is due totheir Lewis and Bronsted acidities. In addition to effects on surfacearea, catalyst modifications (i.e. the modification of alumina with Ce)may also decrease the surface acidity and metal-support interactions ofthe catalyst, thereby enhancing olefin selectivity in oxidative crackingand oxidative dehydrogenation reactions and reducing coke (CO_(x))formation. The catalyst acidity plays a role in metal supportinteractions that affect VO_(x) reducibility. The reducibility mayimpact catalyst activity and selectivity by providing O₂ for oxidationand high acidity not favoring selective oxidation. A number oftechniques have been developed for the characterization of acid-basesurface properties of catalysts. The adsorption of volatile aminesincluding, but not limited to, ammonia (NH₃), pyridine (C₅H₅N),n-butylamine (CH₃CH₂CH₂CH₂NH₂), quinolone (C₉H₇N) and the like is oftenused to determine the acid site concentration of solid catalysts. Theamount of the base remaining on the surface after evacuation isconsidered chemisorbed and serves as a measure of the acid siteconcentration. The adsorbed base concentration as a function ofevacuation temperature can give a site strength distribution. Anothermeans of determining the site strength distribution is calorimetry orthe temperature-programmed desorption (TPD).

Ammonia or NH₃-TPD experiments are used to determine the total acidityof the catalyst. TPD can further give an idea about metal-supportinteractions by modeling NH₃ desorption kinetics and be used todetermine the strength of acid sites available on the catalyst surface.In a preferred embodiment, the catalyst of the present disclosure in anyof its embodiments has a total acidity in the range of 4-16 mL of NH₃per gram of catalyst, preferably 6-15 mL of NH₃ per gram of catalyst,preferably 8-14 mL of NH₃ per gram of catalyst, preferably 10-13 mL ofNH₃ per gram of catalyst, or about 12 mL of NH₃ per gram of catalystwhen measured with a heating rate of 5-20° C./min, preferably 10-15°C./min. In a preferred embodiment, the catalyst of the presentdisclosure has a lower acidity than pure alumina. In a preferredembodiment, the catalyst of the present disclosure has an energy of NH₃desorption established by NH₃-TPD kinetic analysis and an indicator ofactive site metal-support interactions in the range of 1-20 kJ,preferably 2-15 kJ, preferably 4-10 kJ, preferably 5-8 kJ. In certainembodiments, the balance between the acidity and oxygen carryingcapacity of the catalyst may play a role in the oxidative cracking of analkane under gas phase oxygen free conditions. The acidity of thecatalyst may favor cracking; however, excessive cracking producesundesired light/short paraffins (i.e. methane and/or ethane).

According to a second aspect, the present disclosure relates to a methodfor producing the catalyst of the present disclosure in any of itsembodiments, comprising i) mixing an aluminum salt or hydrate with acerium salt or hydrate in a solvent to form an alumina precursorsolution, ii) adding a base to and hydrolyzing the alumina precursorsolution to form the support material comprising alumina modified bycerium, iii) mixing the support material with a solution comprising avanadyl coordination complex or salt in a solvent to form loadedcatalyst precursors, iv) reducing the loaded catalyst precursors with H₂gas to form reduced catalyst precursors, and v) oxidizing the reducedcatalyst precursors with oxygen to form the catalyst.

Two main methods are typically used to prepare supported catalysts. Inthe impregnation method, the solid support or a suspension of the solidsupport is treated with a solution of a precatalyst (for instance ametal salt or metal coordination complex), and the resulting materialthen activated under conditions that will convert the precatalyst to amore active state, such as the metal itself or metal oxides of themetal. In such cases, the catalyst support is usually in the form ofpellets or spheres. Alternatively, supported catalysts can be preparedfrom homogenous solution by co-precipitation. In terms of the presentdisclosure, it is envisaged that the catalyst may be formed by animpregnation method or a co-precipitation method, preferably by animpregnation method, preferably by an impregnation method throughsoaking with an excess solvent. Supports are usually thermally verystable and withstand processes required to activate precatalysts. Forexample, many precatalysts are activated by exposure to a stream ofhydrogen or air (oxygen) at high temperatures, additionally manyprecatalysts may be activated and/or reactivated by oxidation-reductioncycles, again at high temperatures.

In one step of the process, an aluminum salt or hydrate is mixed with acerium salt or hydrate in a solvent to form an alumina precursorsolution. Exemplary aluminum salts or hydrates include, but are notlimited to, aluminum sulfate, aluminum sulfate hydrate, aluminumchloride, aluminum chloride hydrate, aluminum hydroxide, aluminumhydroxide hydrate, aluminum nitrate, aluminum nitrate nonahydrate, andmixtures thereof. Preferably the aluminum salt or hydrate is aluminumnitrate nonahydrate (Al(NO₃)₃.9H₂O). Exemplary cerium salts or hydratesinclude, but are not limited to, cerium sulfate hydrate, cerium sulfate,cerium acetylacetonate, cerium acetylacetonate hydrate, cerium acetate,cerium acetate hydrate, cerium carbonate, cerium carbonate hydrate,cerium oxalate, cerium oxalate hydrate, cerium nitrate hexahydrate, andmixtures thereof. Preferably the cerium salt or hydrate is ceriumnitrate hexahydrate (Ce(NO₃)₃.6H₂O). In a preferred embodiment, thesolvent is polar protic solvent, preferably deionized water as thereaction medium. Exemplary additional polar protic solvents the may beused in addition to, or in lieu of deionized water include, but are notlimited to, methanol, formic acid, n-butanol, isopropanol, n-propanol,ethanol, acetic acid. It is equally envisaged that the reaction may beadapted to be performed in a non-polar solvent (i.e. n-heptane, pentane,cyclopentane, hexane, cyclohexane, benzene, toluene, 1,4-dioxane,chloroform, diethyl ether, and dichloromethane), a polar aprotic solvent(dimethylformamide, tetrahydrofuran, ethyl acetate, acetone,acetonitrile, dimethyl sulfoxide, nitromethane, propylene carbonate),and mixtures thereof.

In a preferred embodiment, the aluminum salt or hydrate is aluminumnitrate nonahydrate (Al(NO₃)₃.9H₂O), the cerium salt or hydrate iscerium nitrate hexahydrate (Ce(NO₃)₃.6H₂O), and the solvent is deionizedwater. In a preferred embodiment the alumina precursor solution has analuminum concentration of 0.5-5.0 M, preferably 0.75-4.0 M, preferably1.0-3.0 M, preferably 1.5-2.5 M, or about 2.0 M. In a preferredembodiment, the weight ratio of the cerium salt or hydrate to thealuminum salt or hydrate is in the range of 1:10000 to 1:100, preferably1:8000 to 1:200, preferably 1:6000 to 1:400, preferably 1:4000 to 1:600,preferably 1:2000 to 1:800, preferably 1:1500 to 1:1000.

In one step of the process, a base is added to hydrolyze the aluminaprecursor solution to form the support material comprising aluminamodified by cerium. As used here in, hydrolysis refers to the cleavageof chemical bonds by the addition of water. Hydrolysis can be thereverse of a condensation reaction; hydrolysis adds water to break down.A common kind of hydrolysis occurs when a salt of a weak acid, a weakbase, or both is dissolved in water which spontaneously ionizes intohydroxide anions and hydronium cations. Additionally, the saltdissociates into its constituent anions and cations. The base may be astrong base (i.e. lithium hydroxide, sodium hydroxide, potassiumhydroxide, etc.) or a weak base (i.e. potassium carbonate, ammoniumcarbonate, ammonium hydroxide, sodium carbonate, calcium carbonate,sodium sulfate), preferably a weak base, most preferably ammoniumcarbonate ((NH₄)CO₃). In a preferred embodiment the base has aconcentration of 0.1-2.5 M, preferably 0.2-2.0 M, preferably 0.5-1.5 M,preferably 0.75-1.25 M, or about 1.0 M. Preferably, the base is addeddropwise and the hydrolysis is performed at a temperature of 20-40° C.,preferably 20-30° C., or about 25° C. for a period of less than 48hours, preferably less than 36 hours, preferably less than 24 hours,preferably less than 18 hours, preferably less than 12 hours, preferablyless than 8 hours, preferably less than 4 hours and optionally withvigorous stirring and/or ultrasonication to achieve a homogeneousmixture.

In a preferred embodiment, the obtained gel support material comprisingalumina modified by cerium is dried before the dispersion of vanadium at20-40° C., preferably 25-35° C. or about 30° C. for a period of up to 48hours, preferably up to 36 hours, preferably up to 24 hours, andfollowing natural drying at a temperature of up to 450° C., preferablyup to 400° C., preferably up to 350° C., preferably up to 300° C.,preferably up to 250° C., preferably up to 200° C., preferably up to175° C., preferably up to 150° C., preferably up to 140° C., preferablyup to 120° C., preferably up to 100° C. for a period of up to 60 hours,preferably up to 48 hours, preferably up to 36 hours, preferably up to24 hours, preferably up to 12 hours, preferably up to 6 hours. In a mostpreferred embodiment, obtained gel support material comprising aluminamodified by cerium is dried at 150° C. for 12 hours, subsequently at200° C. for 12 hours, and finally calcined at 400° C. for 12 hours.

The manner in which the vanadium oxide is deposited onto a support canhave an influence on the properties of the active component in the finalcatalyst. Typically the main method of dispersing vanadium oxide onsupport materials is the classic incipient wetness impregnation methodin a solvent where the vanadium salt is soluble. The impregnation methodis performed by contacting the support with a certain volume of solutioncontaining the dissolved vanadium oxide precursor. If the volume of thesolution is either equal to or less than the pore volume of the support,the technique is referred to as incipient wetness. This particularsynthesis route can show a broad variation of vanadium oxide surfacespecies at all loadings, particularly loadings below monolayer coverage,depending on the synthesis conditions. In one embodiment, this methodmay lead to the formation of crystalline three-dimensional V₂O₅nanoparticles, even at low vanadium oxide loadings. In anotherembodiment, this method may lead to the formation of an amorphousvanadium oxide phase on the surface of the support.

In a preferred embodiment, the loaded catalyst precursors are preparedby an incipient wetness method of impregnation. The alumina-Ce supportcan be immersed in a solution comprising vanadium and/or a vanadium saltor coordination complex. In one embodiment, the vanadium salt orcoordination complex may be a vanadium (IV), vanadium (V) or vanadium(III) salt. Exemplary vanadium salts or coordination complexes include,but are not limited to, ammonium metavanadate in mixtures of water andoxalic acid or methanol and oxalic acid, vanadium (III) acetylacetonate(V(AcAc)₃) or vanadyl acetylacetonate (VO(AcAc)₂) in toluene, VO(iPrO)₃,VO(OC₂H₅)₃, or VO(OC₂H₇)₃ in 2-propanol, as well as vanadyl sulfate,vanadium pentoxide, vanadium (III) chloride, vanadium oxytripropoxide,tetrakis(diethylamido)vanadium(IV), vanadium (IV) chloride, vanadium(III) chloride tetrahydrofuran complex, vanadium (V) oxychloride,vanadium (V) oxyfluoride, and the like. Preferably, the vanadium salt orcoordination complex is vanadium (III) acetylacetonate (V(AcAc)₃) orvanadyl acetylacetonate (VO(AcAc)₂), most preferably vanadylacetylacetonate (VO(AcAc)₂). The vanadium salt is preferably phosphorousfree. In a preferred embodiment, the solvent is a non-polar solvent.Exemplary non-polar solvents include, but are not limited to, pentane,cyclopentane, hexane, cyclohexane, benzene, toluene, 1,4-dioxane,chloroform, diethyl ether, dichloromethane and mixtures thereof,preferably the solvent is toluene. It is equally envisaged that thepresent method may be adapted to incorporate polar protic solventsincluding, but not limited to, formic acid, n-butanol, isopropanol,n-propanol, ethanol, methanol, acetic acid, and water, as well polaraprotic solvents including, but not limited to, tetrahydrofuran, ethylacetate, acetone, dimethylformamide, acetonitrile, dimethyl sulfoxide,nitromethane, propylene carbonate and mixtures thereof.

In a preferred embodiment the vanadium salt is vanadyl acetylacetonateVO(AcAc)₂ and the solvent is toluene. In a preferred embodiment thesolution has a vanadium concentration of 0.01-1.0 M, preferably 0.05-0.5M, preferably 0.1-0.25 M, preferably 0.125-0.2 M, or about 0.15 M. In apreferred embodiment, the weight ratio of vanadyl coordination complexor salt to alumina modified by cerium support is in the range of 1:1 to1:6, preferably 1:1.5 to 1:5, preferably 1:2 to 1:4, or about 1:3.5. Ina preferred embodiment, the mixing of the alumina modified by ceriumsupport material with the vanadyl coordination complex or salt in asolvent is performed at a temperature of 20-40° C., preferably 20-30°C., or about 25° C. for a period of less than 48 hours, preferably lessthan 36 hours, preferably less than 24 hours, preferably less than 18hours, preferably less than 12 hours, preferably less than 10 hours,preferably less than 8 hours and optionally with stirring and/orultrasonication to achieve a homogeneous mixture. In a preferredembodiment, the mixing is performed under vacuum conditions. Aftermixing the solution can be filtered and separated from the solvent toprovided loaded catalyst precursors.

In another embodiment, it is equally envisaged that the method may beadapted to other means of dispersing and depositing the vanadium oxideon the support material. Both adsorption from solution (i.e. grafting)based on attaching vanadia from the solution through reaction withhydroxyl groups on the surface of the support and ion exchange methodspermitting the ionic vanadium oxide species present in an aqueoussolution to be electrostatically attracted by charged sites of thesupport surface have been used. Exemplary other means include, but arenot limited to, vapor-fed flame synthesis, flame spray pyrolysis,sputter deposition, atomic layer deposition and chemical vapordeposition (CVD). For example, chemical vapor deposition (CVD) usesvolatile molecular metal precursors (i.e. O═VCl₃, O═V(OC₂H₅)₃ orO═V(OiPr)₃) to modify oxide support surface and provide a way to controlthe dispersion of the active sites.

In certain embodiments, in addition to the methods employed to dispersevanadium oxide material on different supports, the drying and/orcalcination used for the fixation of the vanadia may be a crucial stepof the catalyst preparation due to the conversion of the initialvanadium species that may result in a broad variety of V_(x)O_(y)species from a nominally simple impregnation process. At highcalcination temperatures, mixed oxide compounds or solid solutions canbe formed with some oxide supports (i.e. AlVO₄). In a preferredembodiment, the loaded catalyst supports or loaded catalyst precursorsare dried before the reduction and the oxidation at room temperature fora period of up to 60 hours, preferably up to 48 hours, preferably up to36 hours, preferably up to 24 hours, and following natural drying beforethe reduction and the oxidation at a temperature of up to 300° C.,preferably up to 250° C., preferably up to 200° C., preferably up to175° C., preferably up to 150° C., preferably up to 140° C., preferablyup to 120° C., preferably up to 100° C. for a period of up to 60 hours,preferably up to 48 hours, preferably up to 36 hours, preferably up to24 hours, preferably up to 12 hours, preferably up to 6 hours.

In one step of the process the loaded catalyst precursors are reducedwith H₂ gas to form reduced catalyst precursors. As used herein,reduction refers to the gain of electrons or a decrease in oxidationstate by a molecule, atom or ion. In a preferred embodiment, the loadedcatalyst precursors are reduced under a flow of hydrogen gas comprising1-40% H₂, preferably 2-20% H₂, preferably 4-18% H₂, preferably 6-16% H₂,preferably 8-14% H₂, or about 10% H₂ as a molar percentage and 60-99%helium or inert gas, preferably 70-95% helium, preferably 80-94% helium,preferably 85-92% helium, or about 90% helium as a molar percentage.Exemplary inert gases include nitrogen (N₂) and argon (Ar), preferablyargon. In a preferred embodiment, the reduction under hydrogen gas flowis performed at a temperature of 300-800° C., preferably 400-780° C.,preferably 500-760° C., preferably 600-750° C., or about 7500° C. for aperiod of 1-18 hours, preferably 2-12 hours, preferably 4-10 hours, orabout 8 hours. In certain embodiments, the reduction of the loadedcatalyst precursors may be performed in a fluidized bed reactor.

In one step of the process the reduced catalyst precursors are oxidizedwith oxygen to form the catalyst of the present disclosure in any of itsembodiments. As used herein, oxidation refers to the loss of electronsor an increase in oxidation state by a molecule, atom or ion. Oxidationreactions are commonly associated with the formation of oxides fromoxygen molecules. Oxygen itself is the most versatile oxidizer. In apreferred embodiment, the reduced catalyst precursors are oxidized underair flow comprising 20-25% O₂, preferably 20.5-22% O₂, or about 21% O₂as a molar percentage and 75-80% N₂, preferably 77-79% N₂, or about 78%N₂ as a molar percentage. In a preferred embodiment, the oxidation underair flow or calcination under air flow is performed at a temperature of300-800° C., preferably 400-780° C., preferably 500-760° C., preferably600-7500° C., or about 7500° C. for a period of 1-18 hours, preferably2-12 hours, preferably 4-10 hours, or about 8 hours.

According to a third aspect, the present disclosure relates to a methodfor the oxidative cracking of an alkane to produce one or more olefinscomprising flowing the alkane through a reactor comprising a catalystchamber loaded with the catalyst of the present disclosure in any of itsembodiments at a temperature in the range of 450-700° C. to form the oneor more olefins and a reduced catalyst.

The general nature of the alkane substrate is not viewed as particularlylimiting to the oxidative cracking and/or oxidative dehydrogenationdescribed herein. As used herein, “alkane” or “paraffin” unlessotherwise specified refers to both branched and straight chain saturatedprimary, secondary and/or tertiary hydrocarbons of typically C₁-C₁₀. Itis equally envisaged that the present disclosure may be adapted tocycloalkanes referring to cyclized alkanes containing one or more ringsand substituted alkanes and/or substituted cycloalkanes referring to atleast one hydrogen atom that is replaced with a non-hydrogen group,provided that normal valencies are maintained and that the substitutionresults in a stable compound. In a preferred embodiment, the alkane isat least one straight-chain linear alkane of C₁ to C₁₀, preferablyC₂-C₈, more preferably C₆ selected from the group consisting of ethane(C₂H₆), propane (C₃H₈), butane (C₄H₁₀, n-butane, isobutane), pentane(C₅H₁₂, n-pentane, isopentane, neopentane), and hexane (C₆H₁₄, n-hexane,isohexane (2-methylpentane, 3-methylpentane, 2,3-dimethylbutane),neohexane) and the one or more olefin is a light olefin selected fromthe group consisting of ethylene, propylene, a butylene (1-butene,(Z)-but-2-ene, (E)-but-2-ene, isobutylene (2-methylpropene)) andbutadiene respectively, more preferably the alkane is hexane, morepreferably n-hexane and the one or more olefins comprise at least one ofethylene, propylene, butylene, and mixtures thereof. In certainembodiments, the alkane (i.e. n-hexane) may be sourced from otherindustrial processes such as those used in the petrochemical industry.Feedstocks generated from petroleum including, but not limited to,ethane, propane, butane, pentane, hexane, naphtha, pet naphtha, pygas,light pygas, and gas oil may serve as substrates for the method ofoxidatively cracking an alkane described herein. In some embodiments,these streams or feedstocks may be processed (i.e. hydroprocessed) priorto the oxidative cracking and/or dehydrogenation. In certainembodiments, the alkane may be hexane and the hexane may be abundantlyavailable from a natural gas source or a refinery off gas source.

As used herein, “cracking” is the process whereby complex organicmolecules such as long chain hydrocarbons are broken down into simplemolecules such as light hydrocarbons, by the breaking of carbon-carbonbonds in the precursors. The rate of cracking and the end products arestrongly dependent on the temperature and presence of catalysts.Cracking is the breakdown of a large alkane into smaller and often moreuseful alkanes and alkenes. Simply, hydrocarbon cracking is the processof breaking a long-chain of hydrocarbons into short ones; the processmay require high temperatures and high pressure. Oxidative cracking isconsidered a promising alternative to the existing thermal/steam/hydrocracking processes to produce olefins with advantages including, but notlimited to, i) being an exothermic oxidation reaction, ii) being anadiabatic process, and iii) having less coke (CO_(x)) formation. Themajor challenges for large scale production of olefins via oxidativecracking of hydrocarbons include thermodynamic favorability of undesiredproducts (lighter alkanes or paraffins by cracking) and completeoxidation of the reactant/products to CO_(x) species. The lattice oxygenpresent in the catalyst of the present disclosure in any of itsembodiments allows the cracking to proceed in an auto-thermal way, wherethe exothermic reaction provides heat for the cracking reaction. In thismechanism (FIG. 1) the lattice oxygen has multiple functions. First, theoxygen influences the radical gas-phase chemistry significantly due tothe type and concentration of chain propagator radicals being greatlyincreased. At higher oxygen partial pressures the radical chemistry isonly slightly influenced by the increasing oxygen concentration.Secondly, the oxygen facilitates the removal of hydrogen from thesurface OH⁻ species that are formed during the activation of alkane onthe catalyst. The catalyst of the present disclosure is used for theoxidative cracking of an alkane, where the catalyst lattice oxygen isutilized for alkane activation. In addition, the catalyst has arelatively high acidity for the cracking reaction and simultaneously,dehydrogenation of alkanes to olefins takes place as shown in FIG. 1.

As used herein, dehydrogenation refers to a chemical reaction thatinvolves the removal of hydrogen from a molecule. It is the reverseprocess of hydrogenation. The dehydrogenation reaction may be conductedon both industrial and laboratory scales. Essentially dehydrogenationconverts saturated materials to unsaturated materials anddehydrogenation processes are used extensively in fine chemicals,oleochemicals, petrochemicals and detergents industries. The mostrelevant industrial pathway in light olefin production is typicallysteam cracking; the alternative fluid catalytic cracking (FCC) is onlyable to produce desired olefins in small concentrations with significantcatalyst deactivation. The FCC catalytic dehydrogenation of alkanes ismore selective but the reaction characteristics pose inherentdifficulties and impose certain technical constraints. For example,thermal dehydrogenation is strongly endothermic and often requiresoperation at both high temperature and high alkane partial pressure. Theoxidative dehydrogenation (ODH) of an alkane, which couples theendothermic dehydrogenation of the alkane with the strongly exothermicoxidation of hydrogen avoids the need for excess internal heat input andconsumes hydrogen. The advantages of the alkane ODH reaction include,but are not limited to, that the reaction is i) exothermic, ii)thermodynamically unrestricted, iii) operates at a much lowertemperature, and iv) minimizes coke (CO_(x)) deposition ensuringlong-term stability of the catalyst.

Under standard operating conditions, an alkane is converted to one ormore light olefins by oxidative cracking of and dehydrogenation in thepresence of the catalyst described herein. The dehydrogenation proceedsin accordance with the chemical equation represented by formula (II),wherein y is a positive whole number, preferably y is 2, 3, or 4, andthe alkane converted is ethane, propane, or butane and the correspondingolefin is ethylene, propylene, or butylene.

C_(y)H_(2y+2)+½V₂O₅→C_(y)H_(2y)+H₂O+½V₂O₃  (II):

In some embodiments the alkane to olefin conversion may be accompaniedby complete oxidation of the alkane or the olefin as side and/orsecondary reactions as represented in formula (III) and formula (IV),wherein y is a positive whole number, preferably y is 2, 3, or 4, morepreferably y is 3, and y is the sum of a and b (y=a+b). The yield ofalkenes or olefins obtained by oxidative dehydrogenation on catalysts islimited by alkene or alkane combustion to carbon oxides CO_(x) (i.e. COand CO₂). In some embodiments a=y and b=0 and CO₂ is the sole combustionproduct considered. The minimization of these undesirable consecutiveand/or parallel combustion reactions is a key in the development ofsuccessful oxidative cracking and/or dehydrogenation catalysts.Additional undesirable side products include the light paraffins(C_(x)H_(2x+2), wherein x<y). For example, if the alkane is hexane,desired products include light olefins (ethylene, propylene, butylene)and undesired products include CO_(x) species (carbon monoxide, carbondioxide) and C1-C5 paraffins (methane, ethane, propane, butane,pentane).

$\begin{matrix}\left. {{C_{y}H_{{2y} + 2}} + {\frac{1}{2}V_{2}O_{5}}}\rightarrow{{a\; {CO}_{2}} + {b\; {CO}} + {\frac{\left( {{2y} + 2} \right)}{2}H_{2}O} + {\frac{1}{2}V_{2}O_{3}}} \right. & ({III}) \\\left. {{C_{y}H_{2y}} + {\frac{1}{2}V_{2}O_{5}}}\rightarrow{{a\; {CO}_{2}} + {b\; {CO}} + {\frac{2y}{2}H_{2}O} + {\frac{1}{2}V_{2}O_{3}}} \right. & ({IV})\end{matrix}$

The performance of the oxidative cracking and/or dehydrogenation can bemodulated by adjusting conditions including, but not limited to,temperature, pressure, reaction time and/or catalyst loading. Oneimportant objective in developing oxidative cracking and/ordehydrogenation catalysts is to reduce the reaction temperature of theprocess to minimize energy consumption. In a preferred embodiment, theoxidative cracking of an alkane to one or more olefin is carried out atemperature in the range of 450-700° C., preferably 450-650° C.,preferably 475-625° C., preferably 500-600° C., preferably 510-580° C.,preferably 520-570° C., preferably 525-560° C., or about 550° C. andpreferably at approximately standard pressure (100 kPa, 1 bar, 14.5 psi,0.9869 atm) such as for example 10-20 psi, preferably 12-18 psi,preferably 14-16 psi, preferably 14.25-15 psi, or approximately14.4-14.8 psi. In a preferred embodiment, the catalyst-alkane feedcontact time is in the range of 1-60 seconds, preferably 1-40 seconds,preferably 2-35 seconds, more preferably 4-30 seconds, preferably 5-25seconds, preferably 10-20 seconds or about 15 seconds. In a preferredembodiment, the catalyst loading or amount of catalyst present in theoxidative cracking and dehydrogenation reaction is in the range of0.05-2.5 g of catalyst per mL of alkane feed injected, preferably0.10-2.4 g/mL, preferably 0.5-2.35 g/mL, preferably 1.0-2.3 g/mL,preferably 1.5-2.25, preferably 1.75-2.2 g/mL, preferably 2.0-2.15 g ofcatalyst per mL of alkane feed injected, or about 2.125 g/mL. Theconditions may vary from these ranges and still provide acceptableconditions for performing the oxidative cracking and/or dehydrogenationof an alkane to one or more olefins utilizing the catalyst of thepresent disclosure.

Oxidative cracking and/or dehydrogenation catalysts are evaluated fortheir percent conversion of the alkane as well as their selectivity to aproduct (i.e. light olefins (ethylene, propylene, butylene), lightparaffins (methane, ethane, propane, butane, pentane) or CO_(x) (COand/or CO₂). The definitions used in calculating the conversion andselectivity are represented for the method of the present disclosureusing the oxidative cracking catalyst are represented in formula (V) andformula (VI) respectively.

$\begin{matrix}{\mspace{79mu} {{{Conversion}\mspace{14mu} {of}\mspace{14mu} {alkane}} = {\frac{{Moles}\mspace{14mu} {of}\mspace{14mu} {alkane}\mspace{14mu} {converted}}{{Moles}\mspace{14mu} {of}\mspace{14mu} {alkane}\mspace{14mu} {fed}} \times 100\%}}} & (V) \\{{{Selectivity}\mspace{14mu} {to}\mspace{14mu} {product}\mspace{14mu} i} = {\frac{{Moles}\mspace{14mu} {of}\mspace{14mu} {product}\mspace{14mu} i}{{Moles}\mspace{14mu} {of}\mspace{14mu} {alkane}\mspace{14mu} {converted}} \times 100\%}} & ({VI})\end{matrix}$

The conversion of alkane (i.e. hexane) (%) can be thought of as moles ofalkane converted divided by moles of alkane fed multiplied by 100% andthe selectivity (i.e. ethylene, propylene, and butylene) to product canbe thought of as moles of product divided by the moles of alkaneconverted multiplied by 100%.

In one embodiment, the method of the present disclosure has an oxidativecracking and dehydrogenation alkane conversion rate as defined withformula (V) of up to 60%, preferably up to 55%, preferably up to 50%/o,preferably up to 45%, preferably up to 40%, preferably up to 35%, suchas for example 5-50%, preferably 10-48%, preferably 15-45%, morepreferably 20-40% and at least 5%, preferably at least 10%, preferablyat least 15%, preferably at least 20%, preferably at least 25%. Inanother embodiment, the alkane is hexane and the method has an alkaneconversion of up to 60%, preferably up to 55%, preferably up to 50%,preferably up to 45%, preferably up to 40%, preferably up to 35%, suchas for example 5-50%, preferably 10-48%, preferably 15-45%, morepreferably 20-40%. In a preferred embodiment, the alkane is hexane andthe one or more olefins comprise ethylene, propylene and butylene andthe method is performed with a catalyst-alkane feed contact time orreaction time of 1-40 seconds, preferably 2-35 seconds, more preferably4-30 seconds, preferably 5-25 seconds, preferably 10-20 at a reactiontemperature of 500-650° C., preferably 505-600° C., preferably 510-580°C., preferably 520-570° C., preferably 525-560° C. and the method has ahexane conversion of 5-50%, preferably 10-48%, preferably 15-45%, morepreferably 20-40%.

In one embodiment, the method of the present disclosure has an oxidativecracking and dehydrogenation light/short olefin (ethylene, propylene,and butylene) selectivity relative to a total percentage of productsformed as defined with formula (VI) of at least 20%, preferably at least25%, preferably at least 30%, preferably at least 35%, preferably atleast 40%, preferably at least 45%, preferably at least 50%, preferablyat least 55%, preferably at least 60% such as for example 20-70%,preferably 30-65%, preferably 35-62%, more preferably 40-60%. In anotherembodiment, the alkane is hexane and the method has an light/shortolefin (ethylene, propylene, and butylene) selectivity relative to atotal percentage of products formed of at least 20%, preferably at least25%, preferably at least 30%, preferably at least 35%, preferably atleast 40%, preferably at least 45%, preferably at least 50%, preferablyat least 55%, preferably at least 60% such as for example 20-70%,preferably 30-65%, preferably 35-62%, more preferably 40-60%. In apreferred embodiment, the alkane is hexane and the one or more olefinscomprises light/short olefins (ethylene, propylene, and butylene) andthe method is performed with a catalyst-alkane feed contact time orreaction time of 1-40 seconds, preferably 2-35 seconds, more preferably4-30 seconds, preferably 5-25 seconds, preferably 10-20 at a reactiontemperature of 500-650° C., preferably 505-600° C., preferably 510-580°C., preferably 520-570° C., preferably 525-560° C. and the method has alight/short olefins selectivity relative to a total percentage ofproducts formed of at least 20%, preferably at least 25%, preferably atleast 30%, preferably at least 35%, preferably at least 40%, preferablyat least 45%, preferably at least 50%, preferably at least 55%,preferably at least 60% such as for example 20-70%, preferably 30-65%,preferably 35-62%, more preferably 40-60%.

In a more preferred embodiment, the alkane is hexane and the one or moreolefins comprises light/short olefins (ethylene, propylene, andbutylene) and the method is performed with a catalyst-alkane feedcontact time or reaction time of 1-40 seconds, preferably 2-35 seconds,more preferably 4-30 seconds, preferably 5-25 seconds, preferably 10-20at a reaction temperature of 500-650° C., preferably 505-600° C.,preferably 510-580° C., preferably 520-570° C., preferably 525-560° C.and the method has a light/short olefins selectivity relative to a totalpercentage of products formed of at least 60%, preferably at least 62%,preferably at least 64%, preferably at least 66%, preferably at least68%, preferably at least 70%, preferably at least 72%, preferably atleast 75%, preferably at least 80% such as for example 60-80%,preferably 62-75%, preferably 64-72%, more preferably 66-70%.

In a preferred embodiment, the method of the present disclosure isperformed with a catalyst-alkane, preferably hexane, feed contact timeor reaction time of 1-40 seconds, preferably 2-35 seconds, morepreferably 4-30 seconds, preferably 5-25 seconds, preferably 10-20 at areaction temperature of 500-650° C., preferably 505-600° C., preferably510-580° C., preferably 520-570° C., preferably 525-560° C. and themethod has a CO_(x) (carbon monoxide, carbon dioxide) or completecombustion selectivity relative to a total percentage of products formedthat is less than the olefin selectivity, and the CO₂ selectivity is nomore than 50%, preferably no more than 45%, preferably no more than 40%,preferably no more than 35%, preferably no more than 30%, preferably nomore than 25%, preferably no more than 20%, preferably no more than 15%,preferably no more than 10% such as for example 5-40%, preferably10-35%, preferably 20-30%.

In a preferred embodiment, the method of the present disclosure isperformed with a catalyst-alkane, preferably hexane, feed contact timeor reaction time of 1-40 seconds, preferably 2-35 seconds, morepreferably 4-30 seconds, preferably 5-25 seconds, preferably 10-20 at areaction temperature of 500-650° C., preferably 505-600° C., preferably510-580° C., preferably 520-570° C., preferably 525-560° C. and themethod has a light/short paraffins (methane, ethane, propane, butane,pentane) selectivity relative to a total percentage of products formedthat is less than the olefin selectivity, and the light/short paraffinsselectivity is no more than 50%, preferably no more than 45%, preferablyno more than 40%, preferably no more than 35%, preferably no more than30%, preferably no more than 25%, preferably no more than 20%,preferably no more than 15%, preferably no more than 10% such as forexample 5-40%, preferably 10-35%, preferably 20-30%.

In a preferred embodiment, the method of the present disclosure andalkane oxidative cracking (OC) and/or alkane oxidative dehydrogenation(ODH) reactions incorporating the catalyst described herein areperformed in a gas phase oxygen-free environment or atmosphere. Thepresence of excess oxygen inside the reactor or catalyst chamberincreases the combustion reaction and therefore CO_(x) production.Preferably, the amount of oxygen available for the reaction iscontrolled by the catalyst available, or lattice oxygen of the catalyst,specifically the vanadium oxide species. By this method, in reducing thecatalyst loading or increasing the alkane feed to catalyst ratio one canfurther minimize the available oxygen and decrease the combustionreaction, thus enhancing light olefin selectivity.

In a preferred embodiment, the reactor is a fluidized bed reactor. Asused herein, a fluidized bed reactor (FBR) is a type of reactor devicethat can be used to carry out a variety of multiphase chemicalreactions. In this type of reactor, a fluid (gas or liquid) is passedthrough a granular solid material (usually a catalyst, preferablyspherically shaped) at high enough velocities to suspend the solid andcause it to behave as though it were a fluid. This process, known asfluidization, imparts many important advantages to the fluidized bedreactor. It is equally envisaged that the method of the presentdisclosure may be adapted to be performed in a fixed-bed reactor, butthis generally results in lower oxidative cracking catalyst activity.

The solid substrate (the catalytic material or catalyst of the presentdisclosure upon which the chemical species react) material in afluidized bed reactor is typically supported by a porous plate known asa distributor, distributor plate or sparger distributor. The fluid isthen forced through the distributor up through the solid material. Atlower fluid velocities, the solids remain in place as the fluid passesthrough the voids in the material. This is referred to as a packed bedreactor. As the fluid velocity is increased, the reactor will reach astage where the force of the fluid on the solids is enough to balancethe weight of the solid material. This stage is referred to as incipientfluidization and occurs at this minimum fluidization velocity. Once thisminimum velocity is surpassed, the contents of the reactor bed begin toexpand and swirl around similar to an agitated tank or boiling pot ofwater. The reactor is now a fluidized bed. Depending on the operatingconditions and properties of the solid phase various flow regimes can beobserved in this type of reactor.

The fluidized bed reactor technology has many advantages including, butnot limited to, uniform particle mixing, uniform temperature gradientsand the ability to operate the reactor in continuous state. Due to theintrinsic fluid-like behavior of the solid material, fluidized beds donot experience poor mixing as in packed beds. The complete mixing allowsfor a uniform product that can often be hard to achieve in other reactordesigns. The elimination of radial and axial concentration gradientsalso allows for better fluid-solid contact, which is essential forreaction efficiency and quality. Many chemical reactions require theaddition or removal of heat. Local hot or cold spots within the reactionbed, often a problem in packed beds, are avoided in fluidized conditionssuch as the fluidized bed reactor. In other reactor types, these localtemperature differences, especially hot spots, can result in productdegradation. Thus fluidized bed reactors are well suited to exothermicreactions. The bed-to-surface heat transfer coefficients for fluidizedbed reactors are also high. The fluidized bed nature of these reactorsallows for the ability to continuously withdraw product and introducenew reactants into the reaction vessel. Operating at a continuousprocess state allows for the more efficient production and removesstartup conditions in batch processes.

In certain embodiments, the fluidizability, reactivity, and stability ofthe catalyst of the present disclosure for experimental laboratory scaleoxidative reactions and/or reaction behaviors may be demonstrated orevaluated in a Plexiglas unit with dimensions matching that of a CRECriser simulator. This type of reactor has a capacity of 50-60 cm³,preferably 51-55 cm³ or about 53 cm³ and is a batch unit designed forcatalyst evaluation and kinetic studies under fluidized bed reactorconditions. The major components of the CREC riser simulator include,but are not limited to, a vacuum box, a series of sampling valves, atimer, two pressure transducers and three temperature controllers. Theproduct gas may be analyzed by gas chromatography (GC) with a thermalconductivity detector (TCD) and flame ionization detector (FID).

The oxidative cracking method of the present disclosure may be performedat various temperatures and contact times. In one embodiment, thecontact times may be chosen to be consistent with catalyst reductiontemperature reported by temperature programmed reduction (TPR) analysis.In a typical procedure, the oxidized catalyst sample of the presentdisclosure is loaded into the reactor basket and the reactor basket ischecked for potential leaks. Following the leak test the system ispurged by flowing pure inert gas, preferably nitrogen or argon, mostpreferably argon. The temperature program is started to heat the reactorto the desired temperature. The inert gas flow is maintained to keep thereactor from any interference of gas phase oxygen. Once the reactorreaches a desired temperature, the inert gas flow is discontinued andthe reactor isolation valve is closed once a desired pressure level isreached. A vacuum pump may be used to evacuate the vacuum box down toless than 100 kPa, preferably less than 50 kPa, preferably less than 25kPa, preferably less than 20 kPa. In one embodiment, the catalyst may befluidized by rotating agitation, preferably by an impeller at a speed of100-5000 rpm, preferably 1000-4500 rpm, preferably 2000-4250 rpm,preferably 3000-4000 rpm. In another embodiment, no agitation (i.e. 0rpm) is necessary to fluidize the catalyst. The alkane feed is injectedinto the reactor using a preloaded gas tight syringe or other means andthe reaction proceeds for a pre-specified amount of time. At thetermination point, the isolation valve between the reactor and vacuumbox may automatically open and transfer all reactant and products to thevacuum box for analysis.

In a preferred embodiment, the method for the oxidative cracking of analkane to one or more olefins utilizing the catalyst of the presentdisclosure in any of its embodiments further comprises i) oxidizing atleast a portion of the reduced catalyst by flowing air through thecatalyst chamber to regenerate the catalyst of the present disclosureand ii) repeating the flowing and the oxidizing at least once with aless than 15% decrease in percent conversion of the alkane, a less than15% decrease in selectivity for the one or more olefins relative to atotal percentage of products formed, or both. In this manner, thecatalyst can be recovered and reused in at least 2 reaction iterations,preferably at least 3, preferably at least 4, preferably at least 5,preferably at least 6, preferably at least 10, preferably at least 15,preferably at least 20 reaction iterations.

The catalyst of the present disclosure can be reformed or regeneratedfrom the reduced catalyst; in this case the regeneration is theoxidation of the reduced vanadium species on the support surface. In apreferred embodiment, the regeneration is oxidation under air flowperformed on the reduced catalyst and is performed at a temperature ofup to 800° C., preferably up to 700° C., preferably up to 600° C.,preferably up to 500° C., preferably up to 400° C. for a period of timeof up to 3 hours, preferably up to 2 hours, preferably up to 1 hour,preferably up to 30 minutes, preferably up to 20 minutes, preferably upto 15 minutes, preferably up to 10 minutes, preferably up to 5 minutes.In a preferred embodiment, the regeneration of the catalyst may beperformed in the catalyst chamber, without the need for flowing of thecatalyst. In one embodiment, the reduced catalyst can flow out of thecatalyst chamber to an additional chamber or re-oxidation chamber, beexposed to air flow to regenerate the catalyst, and flow back tocatalyst chamber for use in subsequent reaction iterations. In apreferred embodiment, catalyst performance remains stable in cycles interms of alkane conversion and light olefin (i.e. ethylene, propylene,butylene and mixtures thereof) selectivity indicating the catalyst'sability to be regenerated which confirms catalyst stability at hightemperatures. In a preferred embodiment, at least a portion of thereduced catalyst is oxidized to regenerate the catalyst per reactioncycle, preferably up to 90%, preferably up to 80%, preferably up to 70%,preferably up to 60%, preferably up to 40%, preferably up to 30%,preferably up to 20%, preferably up to 15%, preferably up to 10% of thereduced catalyst is oxidized to regenerate the catalyst per reactioncycle.

In a preferred embodiment, there is a less than 10% change in percentalkane (i.e. hexane) conversion between the first and second iteration,preferably less than 5%, preferably less than 4%, preferably less than3%, preferably less than 2%, preferably less than a 1% change in percentalkane conversion between the first and second iteration. In anotherembodiment, there is a less than a 20% change in percent alkaneconversion, preferably less than 15%, preferably less than 10%,preferably less than 5%, preferably less than a 2% change in percentalkane conversion between the first and twentieth iteration, preferablybetween the first and fifteenth iteration, preferably between the firstand tenth iteration, preferably between the first and fifth iteration,preferably between the first and fourth iteration, preferably betweenthe first and third iteration, preferably between the first and seconditeration.

In a preferred embodiment, there is a less than 10% change in percentCO_(x) (i.e. carbon monoxide, carbon dioxide, and mixtures thereof)selectivity defined as moles of carbon monoxide and carbon dioxideproduced per moles of alkane (i.e. hexane) converted, relative to atotal percentage of products formed between the first and seconditeration, preferably less than 5%, preferably less than 4%, preferablyless than 3%, preferably less than 2%, preferably less than a 1% changein percent CO_(x) selectivity relative to a total percentage of productsformed between the first and second iteration. In another embodiment,there is a less than a 20% change in percent CO_(x) selectivity relativeto a total percentage of products formed, preferably less than 15%,preferably less than 10%, preferably less than 5%, preferably less than2% change in percent CO_(x) selectivity relative to a total percentageof products formed between the first and twentieth iteration, preferablybetween the first and fifteenth iteration, preferably between the firstand tenth iteration, preferably between the first and fifth iteration,preferably between the first and fourth iteration, preferably betweenthe first and third iteration, preferably between the first and seconditeration.

In a preferred embodiment, there is a less than 10% change in percentlight olefin (i.e. ethylene, propylene, butylene and mixtures thereof)selectivity defined as moles of ethylene, propylene and butyleneproduced per moles of alkane (i.e. hexane) converted, relative to atotal percentage of products formed between the first and seconditeration, preferably less than 5%, preferably less than 4%, preferablyless than 3%, preferably less than 2%, preferably less than a 1% changein percent olefin selectivity relative to a total percentage of productsformed between the first and second iteration. In another embodiment,there is a less than a 20% change in percent olefin selectivity relativeto a total percentage of products formed, preferably less than 15%,preferably less than 10%, preferably less than 5%, preferably less than2% change in percent olefin selectivity relative to a total percentageof products formed between the first and twentieth iteration, preferablybetween the first and fifteenth iteration, preferably between the firstand tenth iteration, preferably between the first and fifth iteration,preferably between the first and fourth iteration, preferably betweenthe first and third iteration, preferably between the first and seconditeration.

The examples below are intended to further illustrate methods protocolsfor preparing and characterizing the catalysts of the presentdisclosure. Further, they are intended to illustrate assessing theproperties and performance of these catalysts. They are not intended tolimit the scope of the claims.

Example 1 Catalyst Synthesis

Aluminum nitrate nona-hydrate, cerium nitrate hexa-hydrate and vanadiumacetylacetonate were obtained from Sigma-Aldrich and used withoutfurther purification. Ammonium carbonate was purchased from FisherLimited and deionized water was used in the preparations of thechemicals. The cerium doped mesoporous γ-Al₂O₃ support was prepared bymodifying the approaches previously presented by X. Shang, et al. and J.Wang, et al. [X. Shang, X. Wang, W. Nie, X. Guo, X. Zou, W. Ding, and X.Lu, “Facile strategy for synthesis of mesoporous crystalline γ-aluminaby partially hydrolyzing aluminum nitrate solution,” J. Mater. Chem.,vol. 22, no. 45, p. 23806, 2012; and J. Wang, K. Shang, Y. Guo, and W.C. Li, “Easy hydrothermal synthesis of external mesoporous γ-Al₂O₃nanorods as excellent supports for Au nanoparticles in CO oxidation,”Microporous Mesoporous Mater., vol. 181, no. 3, pp. 141-145, 2013.—eachincorporated herein by reference in its entirety]. A surfactant freeapproach was used and Ce(NO₃)₃.6H₂O was added to an Al-precursorsolution which was then partially hydrolyzed by ammonium carbonate andsubsequently calcined at 400° C. The crystalline Ce-γ-Al₂O₃ was thusobtained, exhibiting a higher BET surface are, a uniform pore sizedistribution, and a better thermal stability. In this manner,Al(NO₃)₃.9H₂O (37.5 g) was dissolved in deionized water (50 mL), and theappropriate amount of Ce(NO₃)₃.6H₂O was added to the Al-precursorsolution. Subsequently, (NH₄)₂CO₃ (1 molar) was added dropwise to theAl-precursor solution under vigorous stirring, the calcinationtemperature was set at 400° C. for 6 hours. The mixture was thenpartially hydrolyzed by one molar ammonium carbonate solution until theformation of white precipitates was observed. The obtained gel was driedin a oven at 30° C. for 24 hours, and then aged at 150° C. and 200° C.each for a 12 hour interval at a ramping rate of 1° C./min and finallycalcined at 400° C. for 12 hours at a ramping rate of 1° C./min.

Vanadium was dispersed on the Ce-γ-Al₂O₃ using an excessive solventimpregnation method under vacuum conditions. The VO_(x)/Ce-γ-Al₂O₃synthesis method involves five steps: i) active site wet impregnation,ii) filtration, iii) drying, iv) reduction, and v) calcination. Thevanadium content was kept constant at 5 wt % by dissolving 0.86 g ofvanadium acetylacetonate in an excess amount of toluene. The resultantvanadium precursor solution was added dropwise to 3 g of Ce-γ-Al₂O₃solid support under vigorous stirring conditions. The solution wasstirred for 6 hours under vacuum and then filtered. After the filtrationthe solid residue was allowed to dry at room temperature for 24 hoursand then further dried in an oven at 140° C. at a ramping rate of 0.5°C./min for an additional 6 hours. Upon drying the prepared material wasreduced in a fluidizable bed reactor. The source of fluidization andreduction was a 10% H₂ with He gas mixture. The bed temperature of thedried catalyst was raised from the ambient to 750° C. at a ramping rateof 0.5° C./min and held constant at 750° C. for an additional 8 hours.Upon completion of the reduction step, the catalyst was oxidized in anoven at 750° C. using the same ramping rate and time interval employedfor the reduction step.

Example 2 X-Ray Diffraction (XRD) Analysis and Characterization of thePrepared Catalyst

The physio-chemical make-up of the different crystalline phases of thecatalyst were studied by X-ray diffraction (XRD) analysis. The analysiswas conducted using a Rigaku Mini-Flex II bench top XRD diffractometer.Cu—Kα monochromatic radiation (λ=0.15406 nm, 30 kV, 15 mA) over a 10-90°range at a scan rate of 3°/min and with the step size of 0.02 was usedto analyze the diffraction pattern of the prepared catalyst andsupports.

FIG. 2 presents the XRD patterns of the (a) VO_(x)/Ce-γ-Al₂O₃ and (b)Ce—Al₂O₃ support samples. As can be seen, γ-Al₂O₃ peaks appear at 45°and 67° in both Ce-γ-Al₂O₃ support and VO_(x)/Ce-γ-Al₂O₃ patterns [S. A.Al-ghamdi, M. M. Hossain, and H. I. De Lasa, “Kinetic Modeling of EthaneOxidative Dehydrogenation over VO_(x)/Al2O3 Catalyst in a Fluidized-BedRiser Simulator,” 2013; and I. E. Wachs and B. M. Weckhuysen, “Structureand reactivity of surface vanadium oxide species on oxide supports,”Appl. Catal. A Gen., vol. 157, pp. 67-90, 1997.—each incorporated hereinby reference in its entirety]. VO_(x) species including isolated andcrystalline vanadia are present on the support surface. Peaks whichappear at 30° and 37° are relevant to the crystalline VO_(x) phase,which is not favorable for the selectivity of the cracking reaction.This phase can be accounted for by formation of the V—O—V bond andappears at high vanadium loading and at high support acidity [F. Klose,“Selective oxidation of ethane over a VO_(x)/γ-Al₂O₃catalyst—investigation of the reaction network,” Appl. Catal. A Gen.,vol. 260, no. 1, pp. 101-110, March 2004.—incorporated herein byreference in its entirety]. The absence of VO_(x) peaks in the range of5° to 20° is related to either the formation of a highly dispersedamorphous phase of VO_(x) or formation of very small crystals of VO_(x)which are undetectable by the XRD analysis [I. A. Bakare, S. A. Mohamed,S. Al-Ghamdi, S. A. Razzak, M. M. Hossain, and H. I. de Lasa, “Fluidizedbed ODH of ethane to ethylene over VOx-MoOx/γ-Al₂O₃ catalyst: Desorptionkinetics and catalytic activity,” Chem. Eng. J., 2014.—incorporatedherein by reference in its entirety]. Consequently, the amorphous VO_(x)phase indicates good dispersion of VO_(x) species on the supportsurface, which results in controlled oxygen release during oxidativecracking and thus a lower combustion and a higher selectivity toolefins. Furthermore, the VO_(x)/Ce-γ-Al₂O₃ catalyst with 0.2 wt % Cecontent of dopant shows no peaks for CeO₂ and it also the peaks of thesupport structure which is due to the presence of the Ce-γ-Al₂O₃support.

Example 3 Fourier Transform Infrared (FTIR) Spectroscopy Analysis andCharacterization of the Prepared Catalyst

Fourier transform infrared (FTIR) spectra of the prepared samples werecollected using a Nicolet 6700 Thermo Fisher Scientific instrument. Foreach experimental run, 3 mg of the catalyst was thoroughly mixed with400 mg of standard KBr. The excited FTIR spectra were collected over therange of 400-4000 cm⁻¹.

FIG. 3 presents the FTIR spectra of the VO_(x)/Ce-γ-Al₂O₃ catalystsample which shows a broad spectral band between the wavenumber3000-3800 cm⁻¹. The band at 1650 cm⁻¹ corresponds to the O—H stretchingfrequency due to adsorbed water or a surface hydroxyl group. Inaddition, the band at 934 cm⁻¹ is relevant to Al—O—V vibrations and isdue to significant interaction γ-Al₂O₃ which confirms the presence ofisolated VO_(x) species on the surface of the support in agreement withthe XRD patterns [A. M. Elfadly, A. M. Badawi, F. Z. Yehia, Y. A.Mohamed, M. A. Betiha, and A. M. Rabie, “Selective nano aluminasupported vanadium oxide catalysts for oxidative dehydrogenation ofethylbenzene to styrene using CO2 as soft oxidant,” Egypt. J. Pet., vol.22, no. 3, pp. 373-380, December 2013.—incorporated herein by referencein its entirety]. Furthermore, the peaks at 1052 cm⁻¹ and 1100 cm⁻¹ arerelated to the V═O stretching mode and the vibration at 754 cm⁻¹ and1100 cm⁻¹ can be attributed to the V—O—V bond, which indicates thepresence of the VO_(x) crystal phase. Ce—O bond vibrations are reportedto be in the range of 440-500 cm⁻¹. Thus, the FTIR spectra confirms thepresence of isolated and crystalline VO_(x) species on the supportsurface; however, these small crystals of VO_(x) and Ce—O species havenot been detected by XRD.

Example 4

Laser Raman Spectroscopy Analysis and Characterization of the PreparedCatalyst Raman spectra were collected using Yvon Jobin equipment using acooled iHR 320 Horiba spectrometer with a CCD detector, which removesthe elastic laser scattering. The laser source was green type at 532 nmand the laser intensity was 50% at a spectrum window of 50 to 2500. Apowder form of the catalyst sample was used to minimize the possibilityof mass transfer limitations and to ensure that all catalyst particlesin the cell are exposed to the flowing gases.

Raman spectroscopy was used to investigate the nature of the vanadiumoxides on the support surface. FIG. 4 shows the Raman spectra of thecatalyst sample. The appearance of the VO_(x) peaks depends on thevanadium loading and on the surface concentration of the VO_(x) species.Peaks in the range of 200 to 600 cm⁻¹ are from the vibration ofpoly-vanadate and crystalline V₂O₅ which results from the V—O—V bond athigh vanadium loading. Additionally, peaks at 995 and 1130 cm⁻¹ arerelevant to the V═O bond which form isolated VO_(x) species.

The Raman analysis is in agreement with the FTIR results since analysesconfirm the presence of isolated, poly-vanadate and crystalline VO_(x)on the support surface. However, previously published literature reportsthat poly-vanadate and crystalline VO_(x) are formed in small amounts incomparison with isolated VO_(x) species at low vanadium loading and withgood dispersion. In addition, the formation of a surface vanadium oxidelayer on oxide supports is more favorable than crystalline V₂O₅ due tothe surface mobility of vanadium oxide and the lower free surface energyof crystalline V₂O₅ (9-8×10⁻⁶ J cm⁻²) relative to supports(Al₂O₃˜68-70×10⁻⁶ J cm⁻²; ZrO₂˜59-80×10⁻⁶ J cm⁻²; TiO₂˜28-38×10⁻⁶ Jcm⁻²).

Example 5 Temperature Programmed Reduction-Oxidation (TPR/TPO)Characterization of the Prepared Catalyst's Reducibility and Stability

The reducibility and reduction temperature of the oxide catalyst wasdetermined by using the temperature programmed reduction (TPR) analysis.The TPR experiments were carried out on Micrometrics AutoChem II 2920analyzer. For a typical TPR analysis, approximately 200 mg of the freshcatalyst sample was loaded in a quartz U-tube, then the sample washeated to 300° C. for 3 hours in an argon (50 mL/min) environment. Aftercooling to 30° C. in argon (Ar) the reduction was carried out using areducing gas mixture of 90 vol % Ar and 10 vol % H₂, which was heated atthe constant rate of 10° C./min (50 mL/min) up to 900° C. and wasmaintained for 1 hour at 900° C. The H₂ consumption by the oxidecatalyst is recorded using a calibrated thermal conductivity detector(TCD) that measures the change in the concentration of the H₂ in theoutlet stream.

Catalyst activity has been measured using TPR analysis. Activity hasbeen measured in terms of H₂ consumption, which gives an indicationabout the catalyst activity in the oxidative dehydrogenation (ODH)reaction and the oxygen carrying capacity. Catalyst activity in TPR/TPOcycles is similar to reaction/regeneration of the catalysts as expectedduring the actual ODH reaction with ethane. Formula (VII) gives theequation of the TPR reaction and formula (II) gives the equation of thegas phase oxygen free ODH reaction.

V₂O₅+2H₂→V₂O₃+2H₂O  (VII):

C_(y)H_(2y+2)+½V₂O₅→C_(y)H_(2y)+H₂O+½V₂O₃  (II):

In can be seen that in both the TPR reaction (formula (VII)) and the ODHreaction (formula (II)) V₂O₅ is reduced to V₂O₃. In contrast, the TPOcycle (formula (VIII)) represents the catalyst regeneration cyclefollowing the reduction in the TPR reaction. Formula (VIII) gives theequation of the TPO reaction.

V₂O₃+O₂→V₂O₅  (VIII):

FIG. 5 presents the TPR profile of the VO_(x)/Ce-γ-Al₂O catalyst.Furthermore, H₂ consumption was calculated from the TPR profile bycalculating the area under the curve. It was established that the oxidephase is responsible for the overall catalytic activity and selectivityin oxidative dehydrogenation, which increase olefin production in theoxidative cracking reaction. However, the crystalline VO_(x) phase onlyhas a small contribution to the catalyst and a minor effect on thecatalytic activity and selectivity. Moreover, the single peak in the TPRprofile indicates that the V₂O₅ reduction occurs in one step, as informula (VII).

FIG. 5 additionally shows that catalyst stability has been tested usingrepeated TPR/TPO cycles. Repeated TPR/TPO cycles give an indicationregarding catalyst thermal stability and help evaluate alumina phasechange at high temperature as well as the catalysts ability forregeneration (re-oxidization). FIG. 5 shows consistent cyclesdemonstrating good catalyst stability at high temperatures up to 800° C.With regards to the mechanisms of phase transformation of vanadiaspecies, it is reported that these species usually form during thereduction steps, which create oxygen vacancies at the surface. When theconcentration of the oxygen vacancies surpasses a certain criticalvalue, the vacancies aggregate into a vacancy disc, called a shearplane. Part of vanadium oxide may shear so that along the shear planethe linkage between trigonal bipyramids is changed from corner-sharinginto edge-sharing. Thus, another stable structure is formed, which isstoichiometrically different from the original structure [Y. H. Kim andH. Lee, “Redox Property of Vanadium Oxide and Its Behavior in CatalyticOxidation,” vol. 20, no. 12, 1999.—incorporated herein by reference inits entirety].

H₂ consumption was further utilized to calculate the reducible vanadiumamount in the catalyst samples. The percentage of vanadium oxidereduction (oxygen carrying capacity) was calculated using the relationof formula (IX) and formula (X).

$\begin{matrix}{{{fraction}\mspace{14mu} {reduced}\mspace{14mu} \%} = {\frac{W_{V}}{W_{0}} \times 100\%}} & ({IX}) \\{W_{V} = \frac{M\; W_{V} \times V_{H_{2}}}{v \times V_{g}}} & (X)\end{matrix}$

In this formula, (1) W_(V) is the amount of reduced vanadium (g), (2)MW, is the molecular weight of vanadium (g/mol), (3) V_(H2) is thevolume of reacted hydrogen (cm³ at STP), (4) V_(g) is the molar volumeof gas (cm³/mol at STP), (5) W₀ is the initial weight of vanadium (g)and (6) v is the stoichiometric number of hydrogen based on the reactionstoichiometry presented in formula (VII). Assuming that V₂O₅ is theinitial reducible catalyst species on the support, then the reductionreaction equation of formula (VII) applies. Table 1 shows that over therepeated TPR/TPO cycles, the percentage of total VO_(x) speciesreduction was found to be 78-79%. Thus, only approximately 22% of theloaded vanadium is not reducible. These findings agree with the FTIR andRaman results, which confirm the presence of isolated and crystallineVO_(x) species, and as previously described, crystalline VO_(x) speciesdo not contribute to the ODH reaction.

TABLE 1 Reduced vanadium over repeated TPR/TPO cycles (catalyst weight =0.1 g) Total H₂ Reduced consumption vanadium Cycle (cm³/g_(catalyst))(wt %) 1 35 81.0 2 34.5 79.9 3 33.9 78.5 4 33.7 78.0

Example 6

NH₃-Temperature Programmed Desorption (NH₃-TPD) Characterization of thePrepared Catalyst's Acidity

The temperature programmed desorption (TPD) experiments were carried outon the same equipment as the TPR analysis. For each experimental run,approximately 200 mg of the catalyst sample is placed over quartz woolin a quartz U-tube. In a typical procedure, the quartz tube containingthe catalyst is first evacuated under a helium flow rate of 30 mL/minfor 2 hours at 300° C. and then cooled to 100° C. Subsequently, thesamples were saturated with a 4.55% NH₃/He gas mixture at a rate of 50mL/min for 1 hour. The samples were then flushed with the helium (50mL/min) for 1 hour to remove the physically adsorbed NH₃. The TPDanalysis of the ammonia saturated samples was carried out from 100° C.to 750° C. at the heating rate of 10° C./min. The exit concentration ofthe desorbed gases was analyzed by the thermal conductivity detector.

Catalyst total acidity plays a vital role in the reaction of crackinghexane. It is important to optimize catalyst acidity to achieve a highconversion, while at the same time avoiding severe cracking whichproduces methane. Catalyst acidity was estimated using NH₃-TPD analysisusing 0.1 g of the VO_(x)/Ce-γ-Al₂O₃ catalyst sample. FIG. 6 shows NH₃desorption in the temperature range from 120 to 750° C. It can be seenthat NH₃ disported at low and high temperatures which indicates thepresence of strong acid sites. The total volume of desorbed ammonia wasestimated by calculating the area under the curve of the experimentaldata, and found to be 12.2 cm³ NH₃/g of catalyst, which is higher thancomparative exemplary catalysts VO_(x)/c-Al₂O₃ andVO_(x)—MoO_(x)/γ-Al₂O₃ that do not show desorption at high temperature.In addition, peaks have been de-convoluted in order to determine thevolume of ammonia desorbed from weak and strong acid sites. FIG. 6 alsoshows that most of the ammonia was desorbed from weak and medium acidsites and the relevant peak occurs at 208° C. (peak 1), and thecorresponding volume of ammonia was found to be 9.6 cm3/g of catalyst.In contrast, the volume of ammonia that was desorbed at high temperaturewas found to be 2.6 cm3/g of catalyst and the relevant peak occurs at570° C.

As previously described, catalyst acidity has a vital role in oxidativecracking, these acid sites include Bronsted and Lewis acid sites. Thisrelatively high acidity comes from the Ce-γ-Al₂O₃ support surface, whichcontains hydroxyls as confirmed by FTIR analysis (10-15 OH per nm²), thelinear hydroxyls being Bronsted like (H⁺ acceptors) and the bridgedhydroxyls being H⁺ donors [A. Haynes, Concepts of Modern Catalysis andKinetics, vol. 2005, no. 05. 2005, pp. 851-851.—incorporated herein byreference in its entirety]. Furthermore, since alumina was activated ata high enough temperature of 400° C., but not sufficient to cause phasechange, this results in dihydroxylation of Bronsted acid sites whichleads to the formation of Lewis acid sites. The strong acid sitescontribute to the cracking of the feed and products via C—C bondfission, which can follow carbenium or combined ion mechanisms [R. Schl,Concepts in Selective Oxidation of Small Alkane Molecules.2009.—incorporated herein by reference in its entirety].

Example 7

NH₃-Temperature Programmed Desorption (NH₃-TPD) Kinetics Analysis of thePrepared Catalyst

The NH₃-TPD data was further analyzed to estimate desorption kineticparameters (activation energy of desorption and frequency factors).These parameters are important to study catalyst surface characteristicssuch as metal support interactions. The property affects the nature ofVO_(x) on the surface and ultimately catalyst conversion and productselectivity. The model as described in this section was used to estimatethese parameters by setting the following assumptions: (i) homogeneouscatalyst surface, k_(d)=(−E_(des)/RT), (ii) ammonia does not re-adsorbduring the experiment, (iii) uniform adsorbate concentration in the gasflow, and (iv) first order adsorption rate in surface coverage. A highgas flow rate was maintained together with appropriate conditions tosatisfy the previous assumptions, and by performing species balance ondesorbing NH₃ desorption rate can be written as formula (XI).

$\begin{matrix}{r_{des} = {{- {V_{m}\left( \frac{d\; \theta_{des}}{dt} \right)}} = {k_{d,o} \times \theta_{des}^{n}}}} & ({XI})\end{matrix}$

FIG. 6 demonstrates that ammonia desorption follows first order rate [B.Fu, J. Lu, P. C. Stair, G. Xiao, M. C. Kung, and H. H. Kung, “Oxidativedehydrogenation of ethane over alumina-supported Pd catalysts. Effect ofalumina overlayer,” J. Catal., vol. 297, pp. 289-295, January2013.—incorporated herein by reference in its entirety]. Taking thisinto account E_(des) and K_(des,0) can be obtained from formula (XII).

$\begin{matrix}{r_{des} = {{- {V_{m}\left( \frac{d\; \theta_{des}}{dt} \right)}} = {k_{d,0}\theta_{des}{\exp \left\lbrack {{- \frac{E_{des}}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}}} & ({XII})\end{matrix}$

In this formula, (i) θ_(des) is the surface coverage of the adsorbedspecies, (ii) K_(d), k_(d,0) are the desorption constant andpre-exponential factor respectively, (iii) T_(m) is the centeringtemperature in ° C., and by rising the temperature gradually with aconstant value β(one can apply formula (XIII), formula (XIV), formula(XV), formula (XVI) and formula (XVII) resulting in formula (XVIII).

$\begin{matrix}{T = {T_{0} + {\beta \; t}}} & ({XIII}) \\{\frac{dT}{dt} = \beta} & ({XIV}) \\{\frac{d\; \theta_{des}}{dt} = {\frac{d\; \theta_{des}{dT}}{{dT}\mspace{11mu} {dt}} = {\beta \frac{d\; \theta_{des}}{dT}}}} & ({XV}) \\{\frac{d\; \theta_{des}}{dT} = {{- \frac{k_{{des},0}}{\beta \; V_{m}}}\theta_{des}{\exp \left\lbrack {{- \frac{E_{des}}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}} & ({XVI}) \\{\theta_{des} = {1 - \frac{V_{des}}{V_{m}}}} & ({XVII}) \\{\frac{{dV}_{des}}{dT} = {\frac{k_{{des},0}}{\alpha}\left( {1 - \frac{V_{des}}{V_{m}}} \right){\exp \left\lbrack {{- \frac{E_{des}}{R}}\left( {\frac{1}{T} - \frac{1}{T_{m}}} \right)} \right\rbrack}}} & ({XVIII})\end{matrix}$

Formula (XVIII) was fitted to experimental data using MATLAB leastsquare methods at a heating rate (β) taken as 10° C./min. In allexperiments ammonia was pre-adsorbed at 120° C. Table 2 reportsstatistical validations such as confidence interval, degree of freedom,and R² value. FIG. 7 shows the model prediction using formula (XVIII)and the NH₃-TPD experimental data. It is evident that there is goodagreement between the predicted and experimental data, which furtherconfirms the validity of the proposed model. The activation energy forammonia desorption from the VO_(x)/Ce-γ-Al₂O₃ catalyst sample was foundto be 6.5 kJ, which is lower than that which has been reported forcomparative exemplary catalysts. Higher activation energy indicates thatthe V-support interaction is stronger than the V—V interaction andtherefore VO_(x) species isolation has been reduced by formation ofcrystalline VO_(x) as shown in the Raman analysis and FTIR analysis.

TABLE 2 Estimated ammonia-TPD kinetic parameters of the preparedcatalyst sample Parameter Value E_(des) 6.5 (kJ) k_(des) 0.161 (cm3/min)R² 0.998 Degree of freedom 48 Confidence interval 95%

Example 8 Evaluation of the Prepared Catalyst in the Fluidized OxidativeCracking of Hexane

The reactivity and stability of the VO_(x)/Ce-γ-Al₂O₃ catalyst samplewas evaluated using a fluidized CREC (CREC: Chemical Reactor EngineeringCentre) riser simulator, a laboratory scale reactor, capable ofsimulating various sections of a riser/downer reactor. The CREC risersimulator consists of a reactor system connected to a vacuum box througha four port valve. For product analysis, the vacuum box is connected toan online gas chromatograph (GC) by a six port valve. The reactor systemconsists of a chamber holding the catalyst basket, heating elements, animpeller to fluidize the catalyst and a feed injection port.

In the present disclosure, the CREC riser simulator was used to test theviability of a new catalyst for use in the fluidized bed oxidativecracking of hexane given the following: i) close control of catalyst andgas-solid reaction times, providing reaction times similar to thoseexpected in industrial circulating fluidized bed units, ii) shortcontact times with the range of a few seconds, iii) careful control ofthe catalyst/reactant weight ratios and fluidization conditions, and iv)accurate control of reaction conditions including temperature andreactant partial pressures. The oxidative cracking (OC) of hexaneexperiments were conducted with the temperature varied from 525 to 600°C. while the contact timer were adjusted between 5 and 25 seconds. Aftereach reaction run, the catalyst was regenerated by supplied air at 550°C. for 15 minutes. Therefore, the alternative reaction and catalystregeneration were achieved without circulating the catalyst.

In a typical run, 0.85 g of oxidized VO_(x)/Ce-γ-Al₂O₃ catalyst samplewas loaded into the reactor basket and a leak test was conducted.Following the leak test, the system was purged by flowing argon. Thetemperature program was started to heat the reactor to the desiredtemperature. The argon flow was maintained to keep the reactor from anyinterference caused by gas phase oxygen. Once the reactor temperaturereached the desired temperature, the argon flow was discontinued. Thereactor isolation valve was closed when it reached the desired pressurelevel. At this stage the vacuum pump was turned on to evacuate thevacuum box down to 20.7 kPa (3.75 psi). The catalyst was fluidized byrotating the impeller. At this point, the hexane feed (0.4 mL) wasinjected into the reactor by using a preloaded gas tight syringe. Thereaction continued for a pre-specified time. At the termination point,the isolation valve between the reactor and vacuum box openedautomatically and transferred all the reactants and products into thevacuum box. The gas samples in the vacuum box were analyzed using andAgilent 7890A gas chromatograph (GC) equipped with both a thermalconductivity detector (TCD) and a flame ionization detector (FID). Foreach catalytic run, the product samples were analyzed three times toensure the accuracy of the analysis. Finally, the product analysis datawas used to calculate the conversion and selectivity of variousproducts. The definitions used in calculating the conversion andselectivity are represented in formula (V) and formula (VI)respectively.

$\begin{matrix}{\mspace{79mu} {{{Conversion}\mspace{14mu} {of}\mspace{14mu} {hexane}} = {\frac{{Moles}\mspace{14mu} {of}\mspace{14mu} {hexane}\mspace{14mu} {converted}}{{Moles}\mspace{14mu} {of}\mspace{14mu} {hexane}\mspace{14mu} {fed}} \times 100\%}}} & (V) \\{{{Selectivity}\mspace{14mu} {to}\mspace{14mu} {product}\mspace{14mu} i} = {\frac{{Moles}\mspace{14mu} {of}\mspace{14mu} {product}\mspace{14mu} i}{{Moles}\mspace{14mu} {of}\mspace{14mu} {hexane}\mspace{14mu} {converted}} \times 100\%}} & ({VI})\end{matrix}$

FIG. 8 reports the products from the oxidative cracking of n-hexane overa temperature range from 525 to 600° C. It can be seen that anincreasing reaction temperature increases n-hexane conversion due to theformation of various products by cracking and dehydrogenation.Conversion of up to 31% has been obtained at 600° C., and the productsfrom this reaction consist of short olefins (C₂H₄ ethylene, C₃H₆propylene, and C₄H₈ butylene) and paraffins (C₁-C₅) together with CO_(x)gases. In addition, the selectivity to olefins decreases with increasingtemperature (66% to 28%), this can be attributed to the formation of CH₄and CO_(x) by cracking due to acidity and combustion at a high oxygenreleasing rate from the catalyst surface at high temperatures. Theobtained results are in agreement with the TPR and TPD analyses, whichshow that the VO_(x)/Ce-γ-Al₂O₃ catalyst has good activity in terms ofH₂ consumption which indicates catalyst activity in ODH type reactions.With regards to cracking of hexane, the VO_(x)/Ce-γ-Al₂O₃ catalystdemonstrates high acidity (FIG. 6) which contributes to increasingcatalyst conversion. However, high CH₄ and CO_(x) amounts have beenobtained due to side reactions (combustion) and harsh cracking of thefeed and products.

FIG. 9 shows products selectivity and conversion from the oxidativecracking of n-hexane over time at a reaction temperature of 550° C. Itcan be seen that olefins selectivity reaches up to 60% at 15 seconds,which is notably higher than other products (i.e. alkanes, CH₄ andCO_(x)). This data confirms the excellent selectivity to olefins of theVO_(x)/Ce-γ-Al₂O₃ catalyst, which can be attributed to several factors,such as (i) role of the catalyst in dehydrogenation of cracking productsand (ii) controlled release of O₂ from the catalyst surface. Inaddition, it has been reported that the oxygen amount inside the reactoraffects product selectivity severely by increasing combustion products.In oxidative cracking, olefins are produced by cracking anddehydrogenation; thus, the type of O₂ species present also have a largerole in the conversion of saturated hydrocarbons into olefins. In thisregard, the catalyst surface usually contains different types of oxygenspecies (electrophilic and nucleophilic), these O₂ species can be formeddepending on metal support interactions and the type of bonds formed.Furthermore, vanadium has the ability to form different types of bondson an alumina surface such as V—O—V, V═O, and V—O—Al as confirmed by theRaman and FTIR results. The V—O—Al bond has been reported to have themost active and selective oxygen for dehydrogenation, this may offer anexplanation for the observed catalyst selectivity to olefins (FIG. 9).FIG. 9 also shows that the VO_(x)/Ce-γ-Al₂O₃ catalyst demonstrates goodstability during the oxidative cracking reaction in agreement with theTPR/TPO cycles which showed catalyst thermal stability and confirmed thestability of VO_(x) species on the support.

The products from oxidative cracking in a gas phase oxygen freeenvironment are different from those that have been obtained via gasphase oxygen cracking. Previously published literature reports haveshown low selectivity to light olefins (ethylene and propylene) when gasphase oxygen is introduced directly into the reactor. Table 3 shows atypical products distribution in an oxidative cracking reaction run in agas phase oxygen free environment. It can be noted that good selectivityto olefins (C₂ ⁼ to C₄ ⁼) has been obtained (60%) as previouslydescribed. In addition, selectivity to ethylene and propylene is higherthan other products such as butylene and paraffins. This can beattributed to catalyst contribution to the dehydrogenation reaction,through which paraffins are converted into olefins as shown in FIG. 1.However, considerable amounts of CO_(x) and CH₄ are obtained due tocombustion and harsh cracking as previously described at hightemperatures.

TABLE 3 Oxidative cracking of hexane reaction products distribution(reaction time = 15 seconds, hexane feed = 0.4 mL at STP) Reactiontemperature (° C.) 525 550 575 600 Hexane conversion 21 25 29 31 (mole%) Product selectivity (mole %) CH₄ 31 33.5 37 41 C₂H₄ 14.3 15.2 11.49.4 C₃H₆ 10.4 6.14 3.76 2.0 C₄H₈ 8.5 5.7 4.4 3.29 C₂H₆ 1.25 2.0 2.68 3.9C₃H₈ 3.9 1.5 2.08 2.6 C₄H₁₀ 1.59 1.2 0.95 0.82 C₅H₁₂ 0.63 0.42 0.28 0.17CO_(x) 27 30 33 37

Oxidative cracking experiments were further conducted using the reducedform of the catalyst and in the absence of air in order to assess therole of thermal cracking. Pure hexane was injected into the reactor atreaction temperatures of 550 and 575° C. for a reaction time of 15seconds. FIG. 10 shows the reaction products without catalystregeneration. Thus, the only products that were obtained are CO_(x)gases which were produced from combustion and a fraction of CH₄ due tothermal cracking. In addition, the conversion of hexane is almostnegligible in the absence of the catalyst. The absence of surface oxygenwhich is available from VO_(x) on catalyst support severely affects thereaction products formed. This was previously confirmed by the TPR/TPOcycles which show that the presence or reducible vanadium is essentialfor H₂ consumption. Furthermore, previously published literature reportshave presented similar conclusions regarding the VO_(x) reductioneffect. This can be attributed to the nature of VO_(x) species formed onthe surface and the types of bonds present.

The oxide phase is responsible for the overall catalytic activity andselectivity such as VO_(x) species, compared to Nb₂O₅ and β-(Nb,V)₂O₅[E. Heracleous and a Lemonidou, “Ni—Nb—O mixed oxides as highly activeand selective catalysts for ethene production via ethane oxidativedehydrogenation. Part II: Mechanistic aspects and kinetic modeling,” J.Catal., vol. 237, no. 1, pp. 175-189, January 2006.—incorporated hereinby reference in its entirety]. However, the crystalline V₂O₅ phase onlyhas a small contribution to the catalyst and a minor effect on thecatalytic activity and selectivity. The multiple peaks in the TPRanalysis also indicate the presence of several types of oxides inagreement with the XRD and Raman analysis results. With regards to themechanisms of phase transformation of vanadia species, it is reportedthat these species usually form during the reduction steps, which createoxygen vacancies at the surface. When the concentration of the oxygenvacancies surpasses a certain critical value, the vacancies aggregateinto a vacancy disc, called a shear plane. Part of vanadium oxide mayshear so that along the shear plane the linkage between trigonalbipyramids is changed from corner-sharing into edge-sharing. Thus,another stable structure is formed, which is stoichiometricallydifferent from the original structure.

It is of value to compare different types of catalysts for the oxidativecatalytic cracking of n-hexane to investigate the effect of differentreaction conditions, catalyst compositions and reaction mechanisms.Table 4 presents differetypes of catalysts which have been used for theoxidative catalytic cracking of n-hexane including zeolites and metaloxides. Although different reaction setups may have been used, thepresence of gas phase oxygen and the effect of the dehydrogenationreaction are of primary interest. For example, Del-Al-MCM-22 zeolitegives up to 95% conversion, this is largely due to high acidity and asurface area value of 448 m²/g, which is proportional to catalystactivity. However, no dehydrogenation reaction takes place, thus thecracking reaction is the major source of olefins, and consequently onlya 40% olefins selectivity has been achieved. In contrast, thedehydrogenation reaction is reported to accompany the cracking reactionwhen metal oxides such as MoO₂ are used for the hexane cracking. A lowconversion of 2% was obtained, although a higher selectivity to olefinswas achieved (Table 4). This can be attributed to the low acidity of thecatalyst and insufficient acid sites available for cracking.Alternatively, supported metal oxides have been implemented foroxidative cracking, such as Pt/MgAl₂O₄, with the intention that aluminahas high acidity and surface area compared to metal oxides. Highcatalyst activity has been achieved (90% conversion) using Pt/MgAl₂O₄and high olefins selectivity as well (70%); although, this highselectivity was obtained only by introducing H₂ and also at a hightemperature of 850° C. Similarly, gas phase oxidative cracking has beenstudied versus catalytic oxidative cracking over Li/MgO. Boyadjian, etal. have reported that increasing gas phase oxygen decreases selectivityto olefins.

TABLE 4 Comparison of the performance of the prepared VO_(x)/CaO-γ-Al₂O₃(1:1) catalyst with that of other ODH catalysts previously reported inthe literature Conver- Olefin Refer- Catalyst sion selectivity Feed enceDel-Al-MCM-22 40% 90% W/F n-hexane = Yong 6.4-64 g-cat Wang,h/mol-n-hexane, 2015 t = 15 min MoO₂  2% Ethylene 29% n-hexane, Jae HeePropylene 35% pressure 7.3 Song, Butylene 19% Torr 2002 0.1% Pt/MgAl₂O₄90% 70% light C₆/O₂/N₂ = X. Liu, olefins Dec. 15, 1993 et al., (mL/min,STP), 2004 GHSV = 30,000 h 1 Li/MgO 28% 60% light 100 mL/min, Cassiaolefins 10% hexane, 8% oxygen and balance He; WHSV = 15.4 h1VO_(x)/Ce-γ-Al₂O₃ 25% 60% light 0.4 mL hexane, Current olefins 0.85 gcatalyst Disclo- sure

In general, dehydrogenation improves olefins selectivity in oxidativecracking: however, it can be seen that strong acidity and high surfacearea are also major factors that affect the conversion. Additionally, incontrast with previously studied catalysts, no hydrogen supply was needfor the reaction with the VO_(x)/Ce-γ-Al₂O₃ catalyst sample.Furthermore, the type of oxygen involved in the oxidative cracking has alarge effect on olefins selectivity. In the context of gas phase oxygenfree oxidative cracking the source of O₂ is the lattice oxygen of thecatalyst. The lattice oxygen content is present is the form ofnucleophilic (O²⁻, O⁻) and electrophilic (O₂ ⁻) forms and they areselective to different oxidation products. In early stages, the releaseof oxygen favors olefins formation. With increasing reaction time, theoxygen released by the catalyst allows the formation of more CO_(x)products. Therefore, activation of C—H bonds of alkane mainly depends onthe catalyst and oxygen active species present on the surface whicheventually affect the olefin selectivity.

In conclusion, the VO_(x)/Ce-γ-Al₂O₃ catalyst sample for oxidativecracking of hexane to light olefins (i.e. ethylene, propylene, butylene)was investigated. The synthesized catalyst samples were characterizedusing a variety of physiochemical techniques. The gas phase oxygen freeoxidative cracking reactions were performed in a CREC riser simulatorunder various reaction conditions. XRD patterns show peaks of amorphousand crystalline VO_(x), which were confirmed by FTIR and Raman analysisas both V—O—V and V—O—Al bonds were detected representing crystallineand isolated VO_(x) respectively. TPR/TPO oxidation-reduction cyclesdemonstrate good catalyst stability and the VO_(x) species were found tobe active H₂ consumption with 78% reducible vanadium content. NH₃-TPDanalysis shows high acidity which was found to be 12.2 cm³ NH₃/g ofcatalyst and desorption kinetics established using a first orderdesorption model calculate a desorption energy of 6.5 kJ, the lowdesorption energy indicating the presence of the V—V interaction.Reactivity tests in the CREC riser simulator show good hexane conversionof 25% and 66% olefins selectivity at 550° C. reaction temperature and areaction time of 15 seconds. In addition, no gas phase oxygen or H2gases were used in the VO_(x)/Ce-γ-Al₂O₃ catalyst reaction. Oxidativedehydrogenation improved catalyst activity by activation of hexane andincreased olefins selectivity as a result of controlled catalyst latticeoxygen contribution.

Thus, the foregoing discussion discloses and describes merely exemplaryembodiments of the present invention. As will be understood by thoseskilled in the art, the present invention may be embodied in otherspecific forms without departing from the spirit or essentialcharacteristics thereof. Accordingly, the disclosure of the presentinvention is intended to be illustrative, but not limiting of the scopeof the invention, as well as other claims. The disclosure, including anyreadily discernible variants of the teachings herein, defines, in part,the scope of the foregoing claim terminology such that no inventivesubject matter is dedicated to the public.

1.-11. (canceled) 12: A method for the oxidative cracking of an alkaneto produce one or more olefins comprising: flowing the alkane through afluidized bed reactor comprising a catalyst chamber loaded with acatalyst at a temperature in the range of 450-700° C. to form the one ormore olefins and a reduced catalyst, wherein the alkane is a C₂-C₈,straight chain linear alkane, wherein the catalyst comprises: a supportmaterial comprising alumina modified by cerium; and a catalytic materialcomprising one or more vanadium oxides disposed on the support material;wherein the catalyst comprises 1-15% of the one or more selected fromthe group consisting of V₂O₅, VO₂, and V₂O₃ by weight relative to thetotal weight of the catalyst, and wherein the catalyst comprises0.05-1.0% of cerium by weight relative to the total weight of thecatalyst.
 13. (canceled) 14: The method of claim 12, wherein the alkaneis hexane and the one or more olefins comprise at least one of ethylene,propylene, butylene and mixtures thereof. 15: The method of claim 12,further comprising: oxidizing at least a portion of the reduced catalystby flowing air through the catalyst chamber to regenerate the catalyst;and repeating the flowing and the oxidizing at least once with a lessthan 15% decrease in percent conversion of the alkane, a less than 15%decrease in selectivity for the one or more olefins relative to a totalpercentage of products formed, or both. 16: The method of claim 12,wherein the catalyst is present at an amount in the range of 0.50-2.5 gof catalyst per mL of alkane. 17: The method of claim 12, wherein thealkane is hexane and the method has a hexane conversion of 5-50 mol % ata reaction time of 1-40 seconds and a temperature of 500-650° C. 18: Themethod of claim 12, wherein the alkane is hexane and the method has ashort olefins selectivity defined as moles of ethylene, propylene, andbutylene produced per moles of hexane converted of 20-70% at a reactiontime of 1-40 seconds and a temperature of 500-650° C. 19: The method ofclaim 12, wherein the alkane is hexane and the method has a CO₁selectivity defined as moles of carbon monoxide and carbon dioxideproduced per moles of hexane converted of no more than 50% at a reactiontime of 1-40 seconds and a temperature of 500-650° C. 20: The method ofclaim 12, wherein the alkane is hexane and the method has a shortolefins selectivity defined as moles of ethylene, propylene, andbutylene produced per moles of hexane converted of at least 60% at areaction time of 1-40 seconds and a temperature of 500-650° C. 21: Themethod of claim 12, wherein the one or more vanadium oxides form anamorphous phase on the surface of the support material. 22: The methodof claim 12, wherein the one or more vanadium oxides form a crystallinephase on the surface of the support material. 23: The method of claim12, wherein the catalyst comprises at least 50% of V₂O₅ by weightrelative to the total weight of the one or more vanadium oxides. 24: Themethod of claim 12, wherein the catalyst has an average particle size inthe range of 20-160 μm. 25: The method of claim 12, wherein the catalysthas an apparent particle density in the range of 1-10 g/cm³. 26: Themethod of claim 12, wherein the catalyst has a BET surface area in therange of 25-400 m²/g.